Oxidation system with sidedraw secondary reactor

ABSTRACT

Disclosed are process and apparatus for vertical splitting of the oxygen supply to a post-oxidation reactor. Further disclosed are process and apparatus for supplying reaction medium to a post-oxidation reactor at a mid-level inlet. Such apparatus and process can assist in reducing oxygen pinch throughout the post-oxidation reactor.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims priority to the following three U.S. ProvisionalApplication Ser. Nos.: U.S. Provisional Application Ser. No. 61/299,450,filed Jan. 29, 2010, titled “OXIDATION SYSTEM WITH SIDEDRAW SECONDARYREACTOR;” U.S. Provisional Application Ser. No. 61/299,453, filed Jan.29, 2010, titled “OXIDATION SYSTEM WITH SIDEDRAW SECONDARY REACTOR;” andU.S. Provisional Application Ser. No. 61/299,455, filed Jan. 29, 2010,titled “OXIDATION SYSTEM WITH SIDEDRAW SECONDARY REACTOR,” thedisclosures of which are incorporated herein by reference in theirentirety to the extent they do not contradict statements herein.

BACKGROUND

1. Field of the Invention

This invention relates generally to a process for the production of apolycarboxylic acid composition. One aspect of the invention concernsthe partial oxidation of a dialkyl aromatic compound (e.g., para-xylene)to produce a crude aromatic dicarboxylic acid (e.g., crude terephthalicacid), which can thereafter be subjected to purification and separation.Another aspect of the invention concerns an improved reactor system thatprovides for a more effective and economical oxidation process.

2. Description of the Related Art

Liquid-phase oxidation reactions are employed in a variety of existingcommercial processes. For example, liquid-phase oxidation is currentlyused for the oxidation of aldehydes to acids (e.g., propionaldehyde topropionic acid), the oxidation of cyclohexane to adipic acid, and theoxidation of alkyl aromatics to alcohols, acids, or diacids. Aparticularly significant commercial oxidation process in the lattercategory (oxidation of alkyl aromatics) is the liquid-phase catalyticpartial oxidation of para-xylene to terephthalic acid. Terephthalic acidis an important compound with a variety of applications. The primary useof terephthalic acid is as a feedstock in the production of polyethyleneterephthalate (“PET”). PET is a well-known plastic used in greatquantities around the world to make products such as bottles, fibers,and packaging.

In a typical liquid-phase oxidation process, including partial oxidationof para-xylene to terephthalic acid, a liquid-phase feed stream and agas-phase oxidant stream are introduced into a reactor and form amulti-phase reaction medium in the reactor. The liquid-phase feed streamintroduced into the reactor contains at least one oxidizable organiccompound (e.g., para-xylene), while the gas-phase oxidant streamcontains molecular oxygen. At least a portion of the molecular oxygenintroduced into the reactor as a gas dissolves into the liquid phase ofthe reaction medium to provide oxygen availability for the liquid-phasereaction. If the liquid phase of the multi-phase reaction mediumcontains an insufficient concentration of molecular oxygen (i.e., ifcertain portions of the reaction medium are “oxygen-starved”),undesirable side-reactions can generate impurities and/or the intendedreactions can be retarded in rate. If the liquid phase of the reactionmedium contains too little of the oxidizable compound, the rate ofreaction may be undesirably slow. Further, if the liquid phase of thereaction medium contains an excess concentration of the oxidizablecompound, additional undesirable side-reactions can generate impurities.

Conventional liquid-phase oxidation reactors are equipped with agitationmeans for mixing the multi-phase reaction medium contained therein.Agitation of the reaction medium is supplied in an effort to promotedissolution of molecular oxygen into the liquid phase of the reactionmedium, maintain relatively uniform concentrations of dissolved oxygenin the liquid phase of the reaction medium, and maintain relativelyuniform concentrations of the oxidizable organic compound in the liquidphase of the reaction medium.

Agitation of the reaction medium undergoing liquid-phase oxidation isfrequently provided by mechanical agitation means in vessels such as,for example, continuous stirred tank reactors (“CSTRs”). Although CSTRscan provide thorough mixing of the reaction medium, CSTRs have a numberof drawbacks. For example, CSTRs have a relatively high capital cost dueto their requirement for expensive motors, fluid-sealed bearings anddrive shafts, and/or complex stirring mechanisms. Further, the rotatingand/or oscillating mechanical components of conventional CSTRs requireregular maintenance. The labor and shutdown time associated with suchmaintenance adds to the operating cost of CSTRs. However, even withregular maintenance, the mechanical agitation systems employed in CSTRsare prone to mechanical failure and may require replacement overrelatively short periods of time.

Bubble column reactors provide an attractive alternative to CSTRs andother mechanically agitated oxidation reactors. Bubble column reactorsprovide agitation of the reaction medium without requiring expensive andunreliable mechanical equipment. Bubble column reactors typicallyinclude an elongated upright reaction zone within which the reactionmedium is contained. Agitation of the reaction medium in the reactionzone is provided primarily by the natural buoyancy of gas bubbles risingthrough the liquid phase of the reaction medium. This natural-buoyancyagitation provided in bubble column reactors reduces capital andmaintenance costs relative to mechanically agitated reactors. Further,the substantial absence of moving mechanical parts associated withbubble column reactors provides an oxidation system that is less proneto mechanical failure than mechanically agitated reactors.

When liquid-phase partial oxidation of para-xylene is carried out in aconventional oxidation reactor (CSTR or bubble column), the productwithdrawn from the reactor is typically a slurry comprising crudeterephthalic acid (“CTA”) and a mother liquor. CTA contains relativelyhigh levels of impurities (e.g., 4-carboxybenzaldehyde, para-toluicacid, fluorenones, and other color bodies) that render it unsuitable asa feedstock for the production of PET. Thus, the CTA produced inconventional oxidation reactors is typically subjected to a purificationprocess that converts the CTA into purified terephthalic acid (“PTA”)suitable for making PET.

Although advances have been made in the art of liquid-phase oxidationreactions, improvements are still needed.

SUMMARY OF THE INVENTION

One embodiment of the present invention concerns a system for producinga polycarboxylic acid by contacting a slurry with a gas-phase oxidant.The system of this embodiment comprises a primary oxidation reactorcomprising a first slurry outlet and a secondary oxidation reactorcomprising a slurry inlet, a second slurry outlet, a normally loweroxidant inlet, and a normally upper oxidant inlet. In this embodiment,the slurry inlet is in downstream fluid-flow communication the firstslurry outlet, the secondary oxidation reactor defines therein asecondary reaction zone having a maximum length L_(s) and a maximumdiameter D_(s), the normally lower oxidant inlet is spaced from thebottom of the secondary reaction zone by less than 0.5L_(s), thenormally upper oxidant inlet is spaced from the bottom of the secondaryreaction zone by at least 0.5L_(s), and the slurry inlet is spaced fromthe bottom of the secondary reaction zone by a distance in the range offrom about 0.3L_(s) to about 0.9L_(s).

Another embodiment of the present invention concerns a system forproducing a polycarboxylic acid by contacting a slurry produced viaoxidation with a gas-phase oxidant. The system of this embodimentcomprises a primary oxidation reactor comprising a first slurry outletand a secondary oxidation reactor comprising a slurry inlet, a secondslurry outlet, and a normally upper oxidant inlet. In this embodiment,the slurry inlet is in downstream fluid-flow communication with thefirst slurry outlet, the secondary oxidation reactor defines therein asecondary reaction zone having a maximum length L_(s) and a maximumdiameter D_(s), the slurry inlet is spaced from the bottom of thesecondary reaction zone by a distance in the range of from about0.3L_(s) to about 0.9L_(s), and the normally upper oxidant inlet isspaced above the slurry inlet by less than 0.4L_(s).

Still another embodiment of the present invention concerns a method formaking a polycarboxylic acid composition. The method of this embodimentcomprises (a) subjecting a first multi-phase reaction medium comprisingan oxidizable compound to oxidation in a primary reaction zone definedin a primary oxidation reactor to thereby produce a first slurry; and(b) contacting at least a portion of the first slurry with a gas-phaseoxidant in a secondary reaction zone defined in a secondary oxidationreactor to thereby produce a second slurry. In this embodiment, thesecondary reaction zone has a maximum length L_(s) and a maximumdiameter D_(s), a first portion of the gas-phase oxidant is introducedinto the secondary reaction zone at a first oxidant inlet region spacedfrom the bottom of the secondary reaction zone by at least 0.5L_(s),where the first portion of the gas-phase oxidant constitutes in therange of from about 5 to about 49 percent of the total volume of thegas-phase oxidant introduced into the secondary reaction zone, and whereat least a portion of the first slurry is introduced into the secondaryreaction zone at a slurry inlet region spaced from the bottom of thesecondary reaction zone by a distance in the range of from about0.3L_(s) to about 0.9L_(s).

BRIEF DESCRIPTION OF THE DRAWINGS

Embodiments of the invention are described in detail below withreference to the attached drawing figures, wherein:

FIG. 1 is a side view of an oxidation reactor constructed in accordancewith one embodiment of the present invention, particularly illustratingthe introduction of feed, oxidant, and reflux streams into the reactor,the presence of a multi-phase reaction medium in the reactor, and thewithdrawal of a gas and a slurry from the top and bottom of the reactor,respectively;

FIG. 2 is a side view of a bubble column reactor equipped with anexternal secondary oxidation reactor that receives a slurry from asidedraw in the primary oxidation reactor;

FIG. 3 is an expanded sectional bottom view of the sidedraw reactortaken along line 3-3 in FIG. 2, particularly illustrating the locationand configuration of an upper oxidant sparger used to introduce at leasta portion of an oxidant stream into the reactor;

FIG. 4 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating the reactionmedium being theoretically partitioned into 30 horizontal slices ofequal volume in order to quantify certain gradients in the reactionmedium;

FIG. 5 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating first and seconddiscrete 20-percent continuous volumes of the reaction medium that havesubstantially different oxygen concentrations and/or oxygen consumptionrates; and

FIG. 6 is a simplified process flow diagram of a process for making PTAin accordance with an embodiment of the present invention.

DETAILED DESCRIPTION

Various embodiments of the present invention concern the liquid-phasepartial oxidation of an oxidizable compound. Such oxidation can becarried out in the liquid phase of a multi-phase reaction mediumcontained in one or more agitated reactors. Suitable agitated reactorsinclude, for example, bubble-agitated reactors (e.g., bubble columnreactors), mechanically agitated reactors (e.g., continuous stirred tankreactors), and flow agitated reactors (e.g., jet reactors). In one ormore embodiments, the liquid-phase oxidation can be carried out using atleast one bubble column reactor.

As used herein, the term “bubble column reactor” shall denote a reactorfor facilitating chemical reactions in a multi-phase reaction medium,wherein agitation of the reaction medium is provided primarily by theupward movement of gas bubbles through the reaction medium. As usedherein, the term “agitation” shall denote work dissipated into thereaction medium causing fluid flow and/or mixing. As used herein, theterms “majority,” “primarily,” and “predominately” shall mean more than50 percent. As used herein, the term “mechanical agitation” shall denoteagitation of the reaction medium caused by physical movement of a rigidor flexible element(s) against or within the reaction medium. Forexample, mechanical agitation can be provided by rotation, oscillation,and/or vibration of internal stirrers, paddles, vibrators, or acousticaldiaphragms located in the reaction medium. As used herein, the term“flow agitation” shall denote agitation of the reaction medium caused byhigh velocity injection and/or recirculation of one or more fluids inthe reaction medium. For example, flow agitation can be provided bynozzles, ejectors, and/or eductors.

In various embodiments, the portion of the agitation of the reactionmedium in the bubble column reactor during oxidation provided bymechanical and/or flow agitation can be less than about 40 percent, lessthan about 20 percent, or less than 5 percent. Additionally, the amountof mechanical and/or flow agitation imparted to the multi-phase reactionmedium during oxidation can be less than about 3 kilowatts per cubicmeter of the reaction medium, less than about 2 kilowatts per cubicmeter, or less than 1 kilowatt per cubic meter.

Referring now to FIG. 1, a bubble column reactor 20 is illustrated ascomprising a vessel shell 22 having a reaction section 24 and adisengagement section 26. Reaction section 24 defines a reaction zone28, while disengagement section 26 defines a disengagement zone 30. Apredominately liquid-phase feed stream can be introduced into reactionzone 28 via feed inlets 32 a,b,c,d. A predominately gas-phase oxidantstream can be introduced into reaction zone 28 via an oxidant sparger 34located in the lower portion of reaction zone 28. The liquid-phase feedstream and gas-phase oxidant stream cooperatively form a multi-phasereaction medium 36 within reaction zone 28. In various embodiments,multi-phase reaction medium 36 can comprise a liquid phase and a gasphase. In other various embodiments, multiphase reaction medium 36 cancomprise a three-phase medium having solid-phase, liquid-phase, andgas-phase components. The solid-phase component of the reaction medium36 can precipitate within reaction zone 28 as a result of the oxidationreaction carried out in the liquid phase of reaction medium 36. Bubblecolumn reactor 20 includes a slurry outlet 38 located near the bottom ofreaction zone 28 and a gas outlet 40 located near the top ofdisengagement zone 30. A slurry effluent comprising liquid-phase andsolid-phase components of reaction medium 36 can be withdrawn fromreaction zone 28 via slurry outlet 38, while a predominantly gaseouseffluent can be withdrawn from disengagement zone 30 via gas outlet 40.

The liquid-phase feed stream introduced into bubble column reactor 20via feed inlets 32 a,b,c,d can comprise an oxidizable compound, asolvent, and a catalyst system.

The oxidizable compound present in the liquid-phase feed stream cancomprise at least one hydrocarbyl group. In various embodiments, theoxidizable compound can be an aromatic compound. Furthermore, theoxidizable compound can be an aromatic compound with at least oneattached hydrocarbyl group or at least one attached substitutedhydrocarbyl group or at least one attached heteroatom or at least oneattached carboxylic acid function (—COOH). In one or more embodiments,the oxidizable compound can be an aromatic compound with at least oneattached hydrocarbyl group or at least one attached substitutedhydrocarbyl group with each attached group comprising from 1 to 5 carbonatoms. Additionally, the oxidizable compound can be an aromatic compoundhaving exactly two attached groups with each attached group comprisingexactly one carbon atom and consisting of methyl groups and/orsubstituted methyl groups and/or at most one carboxylic acid group.Examples of suitable compounds for use as the oxidizable compoundinclude, but are not limited to, para-xylene, meta-xylene,para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluicacid, and/or acetaldehyde. In one or more embodiments, the oxidizablecompound is para-xylene.

A “hydrocarbyl group,” as defined herein, is at least one carbon atomthat is bonded only to hydrogen atoms or to other carbon atoms. A“substituted hydrocarbyl group,” as defined herein, is at least onecarbon atom bonded to at least one heteroatom and to at least onehydrogen atom. “Heteroatoms,” as defined herein, are all atoms otherthan carbon and hydrogen atoms. Aromatic compounds, as defined herein,comprise an aromatic ring. Such aromatic compounds can have at least 6carbon atoms and, in various embodiments, can have only carbon atoms aspart of the ring. Suitable examples of such aromatic rings include, butare not limited to, benzene, biphenyl, terphenyl, naphthalene, and othercarbon-based fused aromatic rings.

If the oxidizable compound present in the liquid-phase feed stream is anormally-solid compound (i.e., is a solid at standard temperature andpressure), the oxidizable compound can be substantially dissolved in thesolvent when introduced into reaction zone 28. The boiling point of theoxidizable compound at atmospheric pressure can be at least about 50°C., in the range of from about 80 to about 400° C., or in the range offrom 125 to 155° C. The amount of oxidizable compound present in theliquid-phase feed can be in the range of from about 2 to about 40 weightpercent, in the range of from about 4 to about 20 weight percent, or inthe range of from 6 to 15 weight percent.

It is now noted that the oxidizable compound present in the liquid-phasefeed may comprise a combination of two or more different oxidizablechemicals. These two or more different chemical materials can be fedcomingled in the liquid-phase feed stream or may be fed separately inmultiple feed streams. For example, an oxidizable compound comprisingpara-xylene, meta-xylene, para-tolualdehyde, para-toluic acid, andacetaldehyde may be fed to the reactor via a single inlet or multipleseparate inlets.

The solvent present in the liquid-phase feed stream can comprise an acidcomponent and a water component. The solvent can be present in theliquid-phase feed stream at a concentration in the range of from about60 to about 98 weight percent, in the range of from about 80 to about 96weight percent, or in the range of from 85 to 94 weight percent. Theacid component of the solvent can be primarily an organic low molecularweight monocarboxylic acid having 1-6 carbon atoms, or 2 carbon atoms.In various embodiments, the acid component of the solvent can primarilybe acetic acid. The acid component can make up at least about 75 weightpercent of the solvent, at least about 80 weight percent of the solvent,or in the range of from 85 to 98 weight percent of the solvent, with thebalance being water or primarily water. The solvent introduced intobubble column reactor 20 can include small quantities of impurities suchas, for example, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (“4-CBA”), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha-bromo-para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. In various embodiments, the total amount of impurities inthe solvent introduced into bubble column reactor 20 can be less thanabout 3 weight percent.

The catalyst system present in the liquid-phase feed stream can be ahomogeneous, liquid-phase catalyst system capable of promoting oxidation(including partial oxidation) of the oxidizable compound. In variousembodiments, the catalyst system can comprise at least one multivalenttransition metal. In one or more embodiments, the multivalent transitionmetal can comprise cobalt. Additionally, the catalyst system cancomprise cobalt and bromine. Furthermore, the catalyst system cancomprise cobalt, bromine, and manganese.

When cobalt is present in the catalyst system, the amount of cobaltpresent in the liquid-phase feed stream can be such that theconcentration of cobalt in the liquid phase of reaction medium 36 ismaintained in the range of from about 300 to about 6,000 parts permillion by weight (“ppmw”), in the range of from about 700 to about4,200 ppmw, or in the range of from 1,200 to 3,000 ppmw. When bromine ispresent in the catalyst system, the amount of bromine present in theliquid-phase feed stream can be such that the concentration of brominein the liquid phase of reaction medium 36 is maintained in the range offrom about 300 to about 5,000 ppmw, in the range of from about 600 toabout 4,000 ppmw, or in the range of from 900 to 3,000 ppmw. Whenmanganese is present in the catalyst system, the amount of manganesepresent in the liquid-phase feed stream can be such that theconcentration of manganese in the liquid phase of reaction medium 36 ismaintained in the range of from about 20 to about 1,000 ppmw, in therange of from about 40 to about 500 ppmw, or in the range of from 50 to200 ppmw.

The concentrations of the cobalt, bromine, and/or manganese in theliquid phase of reaction medium 36, provided above, are expressed on atime-averaged and volume-averaged basis. As used herein, the term“time-averaged” shall denote an average of at least 10 measurementstaken equally over a continuous period of at least 100 seconds. As usedherein, the term “volume-averaged” shall denote an average of at least10 measurements taken at uniform 3-dimensional spacing throughout acertain volume.

The weight ratio of cobalt to bromine (Co:Br) in the catalyst systemintroduced into reaction zone 28 can be in the range of from about0.25:1 to about 4:1, in the range of from about 0.5:1 to about 3:1, orin the range of from 0.75:1 to 2:1. The weight ratio of cobalt tomanganese (Co:Mn) in the catalyst system introduced into reaction zone28 can be in the range of from about 0.3:1 to about 40:1, in the rangeof from about 5:1 to about 30:1, or in the range of from 10:1 to 25:1.

The liquid-phase feed stream introduced into bubble column reactor 20can include small quantities of impurities such as, for example,toluene, ethylbenzene, para-tolualdehyde, terephthaldehyde, 4-CBA,benzoic acid, para-toluic acid, para-toluic aldehyde,alpha-bromo-para-toluic acid, isophthalic acid, phthalic acid,trimellitic acid, polyaromatics, and/or suspended particulate. Whenbubble column reactor 20 is employed for the production of terephthalicacid, meta-xylene and ortho-xylene are also considered impurities. Invarious embodiments, the total amount of impurities in the liquid-phasefeed stream introduced into bubble column reactor 20 can be less thanabout 3 weight percent.

Although FIG. 1 illustrates an embodiment where the oxidizable compound,the solvent, and the catalyst system are mixed together and introducedinto bubble column reactor 20 as a single feed stream, in an alternativeembodiment, the oxidizable compound, the solvent, and the catalyst canbe separately introduced into bubble column reactor 20. For example, itis possible to feed a pure para-xylene stream into bubble column reactor20 via an inlet separate from the solvent and catalyst inlet(s).

The predominately gas-phase oxidant stream introduced into bubble columnreactor 20 via oxidant sparger 34 comprises molecular oxygen (O₂). Invarious embodiments, the oxidant stream comprises in the range of fromabout 5 to about 40 mole percent molecular oxygen, in the range of fromabout 15 to about 30 mole percent molecular oxygen, or in the range offrom 18 to 24 mole percent molecular oxygen. The balance of the oxidantstream can be comprised primarily of a gas or gasses, such as nitrogen,that are inert to oxidation. In one or more embodiments, the oxidantstream can consist essentially of molecular oxygen and nitrogen. Invarious embodiments, the oxidant stream can be dry air that comprisesabout 21 mole percent molecular oxygen and about 78 to about 81 molepercent nitrogen. In other embodiments, the gas-phase oxidant can beenriched air, and can comprise 25 mole percent, 30 mole percent, 35 molepercent, 40 mole percent, 50 mole percent, 55 mole percent, 60 molepercent, 70 mole percent, or 80 mole percent molecular oxygen. In stillother embodiments, the oxidant stream can comprise substantially pureoxygen.

Referring still to FIG. 1, bubble column reactor 20 can be equipped witha reflux distributor 42 positioned above an upper surface 44 of reactionmedium 36. Reflux distributor 42 is operable to introduce droplets of apredominately liquid-phase reflux stream into disengagement zone 30 byany means of droplet formation known in the art. In various embodiments,reflux distributor 42 can produce a spray of droplets directeddownwardly towards upper surface 44 of reaction medium 36. This downwardspray of droplets can affect (i.e., engage and influence) at least about50 percent, at least about 75 percent, or at least 90 percent of themaximum horizontal cross-sectional area of disengagement zone 30. Thisdownward liquid reflux spray can help prevent foaming at or above uppersurface 44 of reaction medium 36 and can also aid in the disengagementof any liquid or slurry droplets entrained in the upwardly moving gasthat flows towards gas outlet 40. Further, the liquid reflux may serveto reduce the amount of particulates and potentially precipitatingcompounds (e.g., dissolved benzoic acid, para-toluic acid, 4-CBA,terephthalic acid, and catalyst metal salts) exiting in the gaseouseffluent withdrawn from disengagement zone 30 via gas outlet 40. Inaddition, the introduction of reflux droplets into disengagement zone 30can, by a distillation action, be used to adjust the composition of thegaseous effluent withdrawn via gas outlet 40.

The liquid reflux stream introduced into bubble column reactor 20 viareflux distributor 42 can have the same or about the same composition asthe solvent component of the liquid-phase feed stream introduced intobubble column reactor 20 via feed inlets 32 a,b,c,d. Thus, the liquidreflux stream can comprise an acid component and water. The acidcomponent of the reflux stream can be a low molecular weight organicmonocarboxylic acid having 1-6 carbon atoms, or 2 carbon atoms. Invarious embodiments, the acid component of the reflux stream can beacetic acid. Furthermore, the acid component can make up at least about75 weight percent of the reflux stream, at least about 80 weight percentof the reflux stream, or in the range of from 85 to 98 weight percent ofthe reflux stream, with the balance being water or primarily water.Because the reflux stream typically can have the same or substantiallythe same composition as the solvent in the liquid-phase feed stream,when this description refers to the “total solvent” introduced into thereactor, such “total solvent” shall include both the reflux stream andthe solvent portion of the feed stream.

During liquid-phase oxidation in bubble column reactor 20, the feed,oxidant, and reflux streams can be substantially continuously introducedinto reaction zone 28, while the gas and slurry effluent streams aresubstantially continuously withdrawn from reaction zone 28. As usedherein, the term “substantially continuously” shall mean for a period ofat least 10 hours interrupted by less than 10 minutes. During oxidation,the oxidizable compound (e.g., para-xylene) can be substantiallycontinuously introduced into reaction zone 28 at a rate of at leastabout 8,000 kilograms per hour, at a rate in the range of from about15,000 to about 200,000 kilograms per hour, in the range of from about22,000 to about 150,000 kilograms per hour, or in the range of from30,000 to 100,000 kilograms per hour. Although the flow rates of theincoming feed, oxidant, and reflux streams can be substantially steady,it is now noted that one embodiment contemplates pulsing the incomingfeed, oxidant, and/or reflux streams in order to improve mixing and masstransfer. When the incoming feed, oxidant, and/or reflux streams areintroduced in a pulsed fashion, their flow rates can vary within about 0to about 500 percent of the steady-state flow rates recited herein,within about 30 to about 200 percent of the steady-state flow ratesrecited herein, or within 80 to 120 percent of the steady-state flowrates recited herein.

The average space-time rate of reaction (“STR”) in bubble columnoxidation reactor 20 is defined as the mass of the oxidizable compoundfed per unit volume of reaction medium 36 per unit time (e.g., kilogramsof para-xylene fed per cubic meter per hour). In conventional usage, theamount of oxidizable compound not converted to product would typicallybe subtracted from the amount of oxidizable compound in the feed streambefore calculating the STR. However, conversions and yields aretypically high for many of the oxidizable compounds referred to herein(e.g., para-xylene), and it is convenient to define the term herein asstated above. For reasons of capital cost and operating inventory, amongothers, the reaction can be conducted with a high STR. However,conducting the reaction at increasingly higher STR may affect thequality or yield of the partial oxidation. Bubble column reactor 20 maybe particularly useful when the STR of the oxidizable compound (e.g.,para-xylene) is in the range of from about 25 kilograms per cubic meterper hour (“kg/m³/hr.”) to about 400 kg/m³/hr., in the range of fromabout 30 kg/m³/hr. to about 250 kg/m³/hr., in the range of from about 35kg/m³/hr. to about 150 kg/m³/hr., or in the range of from 40 kg/m³/hr.to 100 kg/m³/hr.

The oxygen-STR in bubble column oxidation reactor 20 is defined as theweight of molecular oxygen consumed per unit volume of reaction medium36 per unit time (e.g., kilograms of molecular oxygen consumed per cubicmeter per hour). For reasons of capital cost and oxidative consumptionof solvent, among others, the reaction can be conducted with a highoxygen-STR. However, conducting the reaction at increasingly higheroxygen-STR eventually reduces the quality or yield of the partialoxidation. Without being bound by theory, it appears that this possiblyrelates to the transfer rate of molecular oxygen from the gas phase intothe liquid at the interfacial surface area and thence into the bulkliquid. Too high an oxygen-STR possibly leads to too low a dissolvedoxygen content in the bulk liquid phase of the reaction medium.

The global-average-oxygen-STR is defined herein as the weight of alloxygen consumed in the entire volume of reaction medium 36 per unit time(e.g., kilograms of molecular oxygen consumed per cubic meter per hour).Bubble column reactor 20 may be particularly useful when theglobal-average-oxygen-STR is in the range of from about 25 kg/m³/hr. toabout 400 kg/m³/hr., in the range of from about 30 kg/m³/hr. to about250 kg/m³/hr., in the range of from about 35 kg/m³/hr. to about 150kg/m³/hr., or in the range of from 40 kg/m³/hr. to 100 kg/m³/hr.

During oxidation in bubble column reactor 20, the ratio of the mass flowrate of the total solvent (from both the feed and reflux streams) to themass flow rate of the oxidizable compound entering reaction zone 28 canbe maintained in the range of from about 2:1 to about 50:1, in the rangeof from about 5:1 to about 40:1, or in the range of from 7.5:1 to 25:1.In various embodiments, the ratio of the mass flow rate of solventintroduced as part of the feed stream to the mass flow rate of solventintroduced as part of the reflux stream can be maintained in the rangeof from about 0.5:1 to no reflux stream flow whatsoever, in the range offrom about 0.5:1 to about 4:1, in the range of from about 1:1 to about2:1, or in the range of from 1.25:1 to 1.5:1.

During liquid-phase oxidation in bubble column reactor 20, the oxidantstream can be introduced into bubble column reactor 20 in an amount thatprovides molecular oxygen somewhat exceeding the stoichiometric oxygendemand. The amount of excess molecular oxygen required for best resultswith a particular oxidizable compound affects the overall economics ofthe liquid-phase oxidation. During liquid-phase oxidation in bubblecolumn reactor 20, the ratio of the mass flow rate of the oxidant streamto the mass flow rate of the oxidizable organic compound (e.g.,para-xylene) entering reactor 20 can be maintained in the range of fromabout 0.5:1 to about 20:1, in the range of from about 1:1 to about 10:1,or in the range of from 2:1 to 6:1.

Referring still to FIG. 1, the feed, oxidant, and reflux streamsintroduced into bubble column reactor 20 can cooperatively form at leasta portion of multi-phase reaction medium 36. Reaction medium 36 can be athree-phase medium comprising a solid phase, a liquid phase, and a gasphase. As mentioned above, oxidation of the oxidizable compound (e.g.,para-xylene) can take place predominately in the liquid phase ofreaction medium 36. Thus, the liquid phase of reaction medium 36 cancomprise dissolved oxygen and the oxidizable compound. The exothermicnature of the oxidation reaction that takes place in bubble columnreactor 20 can cause a portion of the solvent (e.g., acetic acid andwater) introduced via feed inlets 32 a,b,c,d to boil/vaporize. Thus, thegas phase of reaction medium 36 in reactor 20 can be formed primarily ofvaporized solvent and an undissolved, unreacted portion of the oxidantstream.

Certain prior art oxidation reactors employ heat exchange tubes/fins toheat or cool the reaction medium. However, such heat exchange structuresmay be undesirable in the inventive reactor and process describedherein. Thus, in various embodiments, bubble column reactor 20 can bedesigned to include substantially no surfaces that contact reactionmedium 36 and exhibit a time-averaged heat flux greater than 30,000watts per meter squared. In addition, in various embodiments, less thanabout 50 percent, less than about 30 percent, or less than 10 percent ofthe time-averaged heat of reaction of reaction medium 36 is be removedby heat exchange surfaces.

The concentration of dissolved oxygen in the liquid phase of reactionmedium 36 is a dynamic balance between the rate of mass transfer fromthe gas phase and the rate of reactive consumption within the liquidphase (i.e., it is not set simply by the partial pressure of molecularoxygen in the supplying gas phase, though this is one factor in thesupply rate of dissolved oxygen and it does affect the limiting upperconcentration of dissolved oxygen). The amount of dissolved oxygenvaries locally, being higher near bubble interfaces. Globally, theamount of dissolved oxygen depends on the balance of supply and demandfactors in different regions of reaction medium 36. Temporally, theamount of dissolved oxygen depends on the uniformity of gas and liquidmixing relative to chemical consumption rates. In designing to matchappropriately the supply of and demand for dissolved oxygen in theliquid phase of reaction medium 36, the time-averaged andvolume-averaged oxygen concentration in the liquid phase of reactionmedium 36 can be maintained above about 1 ppm molar, in the range fromabout 4 to about 1,000 ppm molar, in the range from about 8 to about 500ppm molar, or in the range from 12 to 120 ppm molar.

The liquid-phase oxidation reaction carried out in bubble column reactor20 can be a precipitating reaction that generates solids. In variousembodiments, the liquid-phase oxidation carried out in bubble columnreactor 20 can cause at least about 10 weight percent, at least about 50weight percent, or at least 90 weight percent of the oxidizable compound(e.g., para-xylene) introduced into reaction zone 28 to form a solidcompound (e.g., crude terephthalic acid particles) in reaction medium36. In one or more embodiments, the total amount of solids in reactionmedium 36 can be greater than about 3 weight percent, in the range offrom about 5 to about 40 weight percent, in the range of from about 10to about 35 weight percent, or in the range of from 15 to 30 weightpercent, on a time-averaged and volume-averaged basis. In variousembodiments, a substantial portion of the oxidation product (e.g.,terephthalic acid) produced in bubble column reactor 20 can be presentin reaction medium 36 as solids, as opposed to remaining dissolved inthe liquid phase of reaction medium 36. The amount of the solid phaseoxidation product present in reaction medium 36 can be at least about 25percent by weight of the total oxidation product (solid and liquidphase) in reaction medium 36, at least about 75 percent by weight of thetotal oxidation product in reaction medium 36, or at least 95 percent byweight of the total oxidation product in reaction medium 36. Thenumerical ranges provided above for the amount of solids in reactionmedium 36 apply to substantially steady-state operation of bubble column20 over a substantially continuous period of time, not to start-up,shut-down, or sub-optimal operation of bubble column reactor 20. Theamount of solids in reaction medium 36 is determined by a gravimetricmethod. In this gravimetric method, a representative portion of slurryis withdrawn from the reaction medium and weighed. At conditions thateffectively maintain the overall solid-liquid partitioning presentwithin the reaction medium, free liquid is removed from the solidsportion by sedimentation or filtration, effectively without loss ofprecipitated solids and with less than about 10 percent of the initialliquid mass remaining with the portion of solids. The remaining liquidon the solids is evaporated to dryness, effectively without sublimationof solids. The remaining portion of solids is weighed. The ratio of theweight of the portion of solids to the weight of the original portion ofslurry is the fraction of solids, typically expressed as a percentage.

The precipitating reaction carried out in bubble column reactor 20 cancause fouling (i.e., solids build-up) on the surface of certain rigidstructures that contact reaction medium 36. Thus, in one embodiment,bubble column reactor 20 may be designed to include substantially nointernal heat exchange, stiffing, or baffling structures in reactionzone 28 because such structures would be prone to fouling. If internalstructures are present in reaction zone 28, it is desirable to avoidinternal structures having outer surfaces that include a significantamount of upwardly facing planar surface area because such upwardlyfacing planar surfaces would be highly prone to fouling. Thus, if anyinternal structures are present in reaction zone 28, less than about 20percent of the total upwardly facing exposed outer surface area of suchinternal structures should be formed by substantially planar surfacesinclined less than about 15 degrees from horizontal. Internal structureswith this type of configuration are referred to herein as having a“non-fouling” configuration.

Referring again to FIG. 1, the physical configuration of bubble columnreactor 20 helps provide for optimized oxidation of the oxidizablecompound (e.g., para-xylene) with minimal impurity generation. Invarious embodiments, elongated reaction section 24 of vessel shell 22can include a substantially cylindrical main body 46 and a lower head48. The upper end of reaction zone 28 is defined by a horizontal plane50 extending across the top of cylindrical main body 46. A lower end 52of reaction zone 28 is defined by the lowest internal surface of lowerhead 48. Typically, lower end 52 of reaction zone 28 is locatedproximate the opening for slurry outlet 38. Thus, elongated reactionzone 28 defined within bubble column reactor 20 has a maximum length“L_(p)” measured from the top end 50 to the bottom end 52 of reactionzone 28 along the axis of elongation of cylindrical main body 46. Thelength “L_(p)” of reaction zone 28 can be in the range of from about 10to about 100 meters, in the range of from about 20 to about 75 meters,or in the range of from 25 to 50 meters. Reaction zone 28 has a maximumdiameter (width) “D_(p)” that is typically equal to the maximum internaldiameter of cylindrical main body 46. The maximum diameter D_(p) ofreaction zone 28 can be in the range of from about 1 to about 12 meters,in the range of from about 2 to about 10 meters, in the range of fromabout 3.1 to about 9 meters, or in the range of from 4 to 8 meters. Inone or more embodiments, reaction zone 28 can have a length-to-diameter“L_(p):D_(p)” ratio in the range of from about 6:1 to about 30:1, in therange of from about 8:1 to about 20:1, or in the range of from 9:1 to15:1.

As discussed above, reaction zone 28 of bubble column reactor 20receives multi-phase reaction medium 36. Reaction medium 36 has a bottomend coincident with lower end 52 of reaction zone 28 and a top endlocated at upper surface 44. Upper surface 44 of reaction medium 36 isdefined along a horizontal plane that cuts through reaction zone 28 at avertical location where the contents of reaction zone 28 transitionsfrom a gas-phase-continuous state to a liquid-phase-continuous state.Upper surface 44 can be positioned at the vertical location where thelocal time-averaged gas hold-up of a thin horizontal slice of thecontents of reaction zone 28 is 0.9.

Reaction medium 36 has a maximum height “H_(p)” measured between itsupper and lower ends. The maximum width “W_(p)” of reaction medium 36 istypically equal to the maximum diameter “D_(p)” of cylindrical main body46. During liquid-phase oxidation in bubble column reactor 20, H_(p) canbe maintained at about 60 to about 120 percent of L_(p), about 80 toabout 110 percent of L_(p), or 85 to 100 percent of L_(p). In variousembodiments, reaction medium 36 can have a height-to-width “H_(p):W_(p)”ratio greater than about 3:1, in the range of from about 7:1 to about25:1, in the range of from about 8:1 to about 20:1, or in the range offrom 9:1 to 15:1. In one embodiment of the invention, L_(p)=H_(p) andD_(p)=W_(p) so that various dimensions or ratios provide herein forL_(p) and D_(p) also apply to H_(p) and W_(p), and vice-versa.

The relatively high L_(p):D_(p) and H_(p):W_(p) ratios provided inaccordance with an embodiment of the invention can contribute to severalimportant advantages of the inventive system. As discussed in furtherdetail below, it has been discovered that higher L_(p):D_(p) andH_(p):W_(p) ratios, as well as certain other features discussed below,can promote beneficial vertical gradients in the concentrations ofmolecular oxygen and/or the oxidizable compound (e.g., para-xylene) inreaction medium 36. Contrary to conventional wisdom, which would favor awell-mixed reaction medium with relatively uniform concentrationsthroughout, it has been discovered that the vertical staging of theoxygen and/or the oxidizable compound concentrations facilitate a moreeffective and economical oxidation reaction. Minimizing the oxygen andoxidizable compound concentrations near the top of reaction medium 36can help avoid loss of unreacted oxygen and unreacted oxidizablecompound through upper gas outlet 40. However, if the concentrations ofoxidizable compound and unreacted oxygen are low throughout reactionmedium 36, then the rate and/or selectivity of oxidation are reduced.Thus, in various embodiments, the concentrations of molecular oxygenand/or the oxidizable compound can be significantly higher near thebottom of reaction medium 36 than near the top of reaction medium 36.

In addition, high L_(p):D_(p) and H_(p):W_(p) ratios can cause thepressure at the bottom of reaction medium 36 to be substantially greaterthan the pressure at the top of reaction medium 36. This verticalpressure gradient is a result of the height and density of reactionmedium 36. One advantage of this vertical pressure gradient is that theelevated pressure at the bottom of the vessel drives more oxygensolubility and mass transfer than would otherwise be achievable atcomparable temperatures and overhead pressures in shallow reactors.Thus, the oxidation reaction can be carried out at lower temperaturesthan would be required in a shallower vessel. When bubble column reactor20 is used for the partial oxidation of para-xylene to crudeterephthalic acid (CTA), the ability to operate at lower reactiontemperatures with the same or better oxygen mass transfer rates has anumber of advantages. For example, low temperature oxidation ofpara-xylene reduces the amount of solvent burned during the reaction. Asdiscussed in further detail below, low temperature oxidation also favorsthe formation of small, high surface area, loosely bound, easilydissolved CTA particles, which can be subjected to more economicalpurification techniques than the large, low surface area, dense CTAparticles produced by conventional high temperature oxidation processes.

During oxidation in reactor 20, the time-averaged and volume-averagedtemperature of reaction medium 36 can be maintained in the range of fromabout 125 to about 200° C., in the range of from about 140 to about 180°C., or in the range of from 150 to 170° C. The overhead pressure abovereaction medium 36 can be maintained in the range of from about 1 toabout 20 bar gauge (“barg”), in the range of from about 2 to about 12barg, or in the range of from 4 to 8 barg. The pressure differencebetween the top of reaction medium 36 and the bottom of reaction medium36 can be in the range of from about 0.4 to about 5 bar, in the range offrom about 0.7 to about 3 bar, or in the range of from 1 to 2 bar.Although the overhead pressure above reaction medium 36 can generally bemaintained at a relatively constant value, one embodiment contemplatespulsing the overhead pressure to facilitate improved mixing and/or masstransfer in reaction medium 36. When the overhead pressure is pulsed,the pulsed pressures can range between about 60 to about 140 percent,between about 85 and about 115 percent, or between 95 and 105 percent ofthe steady-state overhead pressure recited herein.

A further advantage of the high L_(p):D_(p) ratio of reaction zone 28 isthat it can contribute to an increase in the average superficialvelocity of reaction medium 36. The term “superficial velocity” and“superficial gas velocity,” as used herein with reference to reactionmedium 36, shall denote the volumetric flow rate of the gas phase ofreaction medium 36 at an elevation in the reactor divided by thehorizontal cross-sectional area of the reactor at that elevation. Theincreased superficial velocity provided by the high L_(p):D_(p) ratio ofreaction zone 28 can promote local mixing and increase the gas hold-upof reaction medium 36. The time-averaged superficial velocities ofreaction medium 36 at one-quarter height, half height, and/orthree-quarter height of reaction medium 36 can be greater than about 0.3meters per second, in the range of from about 0.8 to about 5 meters persecond, in the range of from about 0.9 to about 4 meters per second, orin the range of from 1 to 3 meters per second.

Referring still to FIG. 1, disengagement section 26 of bubble columnreactor 20 can simply be a widened portion of vessel shell 22 locatedimmediately above reaction section 24. Disengagement section 26 reducesthe velocity of the upwardly-flowing gas phase in bubble column reactor20 as the gas phase rises above the upper surface 44 of reaction medium36 and approaches gas outlet 40. This reduction in the upward velocityof the gas phase helps facilitate removal of entrained liquids and/orsolids in the upwardly flowing gas phase and thereby reduces undesirableloss of certain components present in the liquid phase of reactionmedium 36.

Disengagement section 26 can include a generally frustoconicaltransition wall 54, a generally cylindrical broad sidewall 56, and anupper head 58. The narrow lower end of transition wall 54 is coupled tothe top of cylindrical main body 46 of reaction section 24. The wideupper end of transition wall 54 is coupled to the bottom of broadsidewall 56. Transition wall 54 can extend upwardly and outwardly fromits narrow lower end at an angle in the range of from about 10 to about70 degrees from vertical, in the range of about 15 to about 50 degreesfrom vertical, or in the range of from 15 to 45 degrees from vertical.Broad sidewall 56 has a maximum diameter “X” that is generally greaterthan the maximum diameter D_(p) of reaction section 24, though when theupper portion of reaction section 24 has a smaller diameter than theoverall maximum diameter of reaction section 24, then X may actually besmaller than D_(p). In various embodiments, the ratio of the diameter ofbroad sidewall 56 to the maximum diameter of reaction section 24“X:D_(p)” can be in the range of from about 0.8:1 to about 4:1, or inthe range of from 1.1:1 to 2:1. Upper head 58 is coupled to the top ofbroad sidewall 56. Upper head 58 can be a generally elliptical headmember defining a central opening that permits gas to escapedisengagement zone 30 via gas outlet 40. Alternatively, upper head 58may be of any shape, including conical. Disengagement zone 30 has amaximum height “Y” measured from the top 50 of reaction zone 28 to theupper-most portion of disengagement zone 30. The ratio of the length ofreaction zone 28 to the height of disengagement zone 30 “L_(p):Y” can bein the range of from about 2:1 to about 24:1, in the range of from about3:1 to about 20:1, or in the range of from 4:1 to 16:1.

Referring still to FIG. 1, during operation a gas-phase oxidant (e.g.,air) can be introduced into reaction zone 28 via oxidant inlets 66 a,band oxidant sparger 34. Oxidant sparger 34 can have any shape orconfiguration that permits passage of the gas-phase oxidant intoreaction zone 28. For instance, oxidant sparger 34 can comprise acircular or polygonal (e.g., octagonal) ring member defining a pluralityof oxidant discharge openings. In various embodiments, some or all ofthe oxidant discharge openings can be configured to discharge thegas-phase oxidant in a generally downward direction. Regardless of thespecific configuration of oxidant sparger 34, the oxidant sparger can bephysically configured and operated in a manner that minimizes thepressure drop associated with discharging the oxidant stream through theoxidant discharge openings and into the reaction zone. Such pressuredrop is calculated as the time-averaged static pressure of the oxidantstream inside the flow conduit at oxidant inlets 66 a,b of the oxidantsparger minus the time-averaged static pressure in the reaction zone atthe elevation where one-half of the oxidant stream is introduced abovethat vertical location and one-half of the oxidant stream is introducedbelow that vertical location. In various embodiments, the time-averagedpressure drop associated with discharging the oxidant stream from theoxidant sparger 34 can be less than about 0.3 megapascal (“MPa”), lessthan about 0.2 MPa, less than about 0.1 MPa, or less than 0.05 MPa.

Optionally, a continuous or intermittent flush can be provided tooxidant sparger 34 with a liquid (e.g., acetic acid, water, and/orpara-xylene) to prevent fouling of the oxidant sparger with solids. Whensuch a liquid flush is employed, an effective amount of the liquid(i.e., not just the minor amount of liquid droplets that might naturallybe present in the oxidant stream) can be passed through the oxidantsparger and out of the oxidant openings for at least one period of morethan one minute each day. When a liquid is continuously or periodicallydischarged from oxidant sparger 34, the time-averaged ratio of the massflow rate of the liquid through the oxidant sparger to the mass flowrate of the molecular oxygen through the oxidant sparger can be in therange of from about 0.05:1 to about 30:1, in the range of from about0.1:1 to about 2:1, or in the range of from 0.2:1 to 1:1.

In many conventional bubble column reactors containing a multi-phasereaction medium, substantially all of the reaction medium located belowthe oxidant sparger (or other mechanism for introducing the oxidantstream into the reaction zone) has a very low gas hold-up value. Asknown in the art, “gas hold-up” is simply the volume fraction of amulti-phase medium that is in the gaseous state. Zones of low gashold-up in a medium can also be referred to as “unaerated” zones. Inmany conventional slurry bubble column reactors, a significant portionof the total volume of the reaction medium is located below the oxidantsparger (or other mechanism for introducing the oxidant stream into thereaction zone). Thus, a significant portion of the reaction mediumpresent at the bottom of conventional bubble column reactors isunaerated.

It has been discovered that minimizing the amount of unaerated zones ina reaction medium subjected to oxidization in a bubble column reactorcan minimize the generation of certain types of undesirable impurities.Unaerated zones of a reaction medium contain relatively few oxidantbubbles. This low volume of oxidant bubbles reduces the amount ofmolecular oxygen available for dissolution into the liquid phase of thereaction medium. Thus, the liquid phase in an unaerated zone of thereaction medium has a relatively low concentration of molecular oxygen.These oxygen-starved, unaerated zones of the reaction medium have atendency to promote undesirable side reactions, rather than the desiredoxidation reaction. For example, when para-xylene is partially oxidizedto form terephthalic acid, insufficient oxygen availability in theliquid phase of the reaction medium can cause the formation ofundesirably high quantities of benzoic acid and coupled aromatic rings,notably including highly undesirable colored molecules known asfluorenones and anthraquinones.

In accordance one or more embodiments, liquid-phase oxidation can becarried out in a bubble column reactor configured and operated in amanner such that the volume fraction of the reaction medium with low gashold-up values is minimized. This minimization of unaerated zones can bequantified by theoretically partitioning the entire volume of thereaction medium into 2,000 discrete horizontal slices of uniform volume.With the exception of the highest and lowest horizontal slices, eachhorizontal slice is a discrete volume bounded on its sides by thesidewall of the reactor and bounded on its top and bottom by imaginaryhorizontal planes. The highest horizontal slice is bounded on its bottomby an imaginary horizontal plane and on its top by the upper surface ofthe reaction medium. The lowest horizontal slice is bounded on its topby an imaginary horizontal plane and on its bottom by the lower end ofthe vessel. Once the reaction medium has been theoretically partitionedinto 2,000 discrete horizontal slices of equal volume, the time-averagedand volume-averaged gas hold-up of each horizontal slice can bedetermined. When this method of quantifying the amount of unaeratedzones is employed, the number of horizontal slices having atime-averaged and volume-averaged gas hold-up less than 0.1 can be lessthan 30, less than 15, less than 6, less than 4, or less than 2.Furthermore, the number of horizontal slices having a gas hold-up lessthan 0.2 can be less than 80, less than 40, less than 20, less than 12,or less than 5. Also, the number of horizontal slices having a gashold-up less than 0.3 can be less than 120, less than 80, less than 40,less than 20, or less than 15.

Referring still to FIG. 1, it has been discovered that positioningoxidant sparger 34 lower in reaction zone 28 provides severaladvantages, including reduction of the amount of unaerated zones inreaction medium 36. Given a height “H_(p)” of reaction medium 36, alength “L_(p)” of reaction zone 28, and a maximum diameter “D_(p)” ofreaction zone 28, a majority of the oxidant stream can be introducedinto reaction zone 28 within about 0.025H_(p), 0.022L_(p), and/or0.25D_(p) of lower end 52 of reaction zone 28, within about 0.02H_(p),0.018L_(p), and/or 0.2D_(p) of lower end 52 of reaction zone 28, orwithin 0.015H_(p), 0.013L_(p), and/or 0.15D_(p) of lower end 52 ofreaction zone 28.

In addition to the advantages provided by minimizing unaerated zones(i.e., zones with low gas hold-up) in reaction medium 36, it has beendiscovered that oxidation can be enhanced by maximizing the gas hold-upof the entire reaction medium 36. Reaction medium 36 can have atime-averaged and volume-averaged gas hold-up of at least about 0.4, inthe range of from about 0.6 to about 0.9, or in the range of from 0.65to 0.85. Several physical and operational attributes of bubble columnreactor 20 contribute to the high gas hold-up discussed above. Forexample, for a given reactor size and flow of oxidant stream, the highL_(p):D_(p) ratio of reaction zone 28 yields a lower diameter whichincreases the superficial velocity in reaction medium 36 which in turnincreases gas hold-up. Additionally, the actual diameter of a bubblecolumn and the L_(p):D_(p) ratio are known to influence the average gashold-up even for a given constant superficial velocity. In addition, theminimization of unaerated zones, particularly in the bottom of reactionzone 28, contributes to an increased gas hold-up value. Further, theoverhead pressure and mechanical configuration of the bubble columnreactor can affect operating stability at the high superficialvelocities and gas hold-up values disclosed herein.

Referring still to FIG. 1, it has been discovered that improveddistribution of the oxidizable compound (e.g., para-xylene) in reactionmedium 36 can be provided by introducing the liquid-phase feed streaminto reaction zone 28 at multiple vertically-spaced locations. Invarious embodiments, the liquid-phase feed stream can be introduced intoreaction zone 28 via at least 3 feed openings, or at least 4 feedopenings. As used herein, the term “feed openings” shall denote openingswhere the liquid-phase feed stream is discharged into reaction zone 28for mixing with reaction medium 36. In one or more embodiments, at least2 of the feed openings can be vertically-spaced from one another by atleast about 0.5D_(p), at least about 1.5D_(p), or at least 3D_(p).However, the highest feed opening can be vertically-spaced from thelowest oxidant opening by not more than about 0.75H_(p), 0.65L_(p),and/or 8D_(p); not more than about 0.5H_(p), 0.4L_(p), and/or 5D_(p); ornot more than 0.4H_(p), 0.35L_(p), and/or 4D_(p).

Although it is desirable to introduce the liquid-phase feed stream atmultiple vertical locations, it has also been discovered that improveddistribution of the oxidizable compound in reaction medium 36 isprovided if the majority of the liquid-phase feed stream is introducedinto the bottom half of reaction medium 36 and/or reaction zone 28. Invarious embodiments, at least about 75 weight percent or at least 90weight percent of the liquid-phase feed stream can be introduced intothe bottom half of reaction medium 36 and/or reaction zone 28. Inaddition, at least about 30 weight percent of the liquid-phase feedstream can be introduced into reaction zone 28 within about 1.5D_(p) ofthe lowest vertical location where the oxidant stream is introduced intoreaction zone 28. This lowest vertical location where the oxidant streamis introduced into reaction zone 28 is typically at the bottom ofoxidant sparger 34; however, a variety of alternative configurations forintroducing the oxidant stream into reaction zone 28 are contemplated byvarious embodiments. In one or more embodiments, at least about 50weight percent of the liquid-phase feed can be introduced within about2.5D_(p) of the lowest vertical location where the oxidant stream isintroduced into reaction zone 28. In other embodiments, at least about75 weight percent of the liquid-phase feed stream can be introducedwithin about 5D_(p) of the lowest vertical location where the oxidantstream is introduced into reaction zone 28.

Each feed opening defines an open area through which the feed isdischarged. In various embodiments, at least about 30 percent of thecumulative open area of all the feed inlets can be located within about1.5D_(p) of the lowest vertical location where the oxidant stream isintroduced into reaction zone 28. In other embodiments, at least about50 percent of the cumulative open area of all the feed inlets can belocated within about 2.5D_(p) of the lowest vertical location where theoxidant stream is introduced into reaction zone 28. In still otherembodiments, at least about 75 percent of the cumulative open area ofall the feed inlets can be located within about 5D_(p) of the lowestvertical location where the oxidant stream is introduced into reactionzone 28.

Referring still to FIG. 1, in one or more embodiments, feed inlets 32a,b,c,d can simply be a series of vertically-aligned openings along oneside of vessel shell 22. These feed openings can have substantiallysimilar diameters of less than about 7 centimeters, in the range of fromabout 0.25 to about 5 centimeters, or in the range of from 0.4 to 2centimeters. Bubble column reactor 20 can be equipped with a system forcontrolling the flow rate of the liquid-phase feed stream out of eachfeed opening. Such flow control system can include an individual flowcontrol valve 74 a,b,c,d for each respective feed inlet 32 a,b,c,d. Inaddition, bubble column reactor 20 can be equipped with a flow controlsystem that allows at least a portion of the liquid-phase feed stream tobe introduced into reaction zone 28 at an elevated inlet superficialvelocity of at least about 2 meters per second, at least about 5 metersper second, at least about 6 meters per second, or in the range of from8 to 20 meters per second. As used herein, the term “inlet superficialvelocity” denotes the time-averaged volumetric flow rate of the feedstream out of the feed opening divided by the area of the feed opening.In various embodiments, at least about 50 weight percent of the feedstream can be introduced into reaction zone 28 at an elevated inletsuperficial velocity. In one or more embodiments, substantially all thefeed stream is introduced into reaction zone 28 at an elevated inletsuperficial velocity.

Referring now to FIG. 2, there is illustrated a reactor system 100comprising a primary oxidation reactor 102 and a secondary oxidationreactor 104. Primary oxidation reactor 102 can be configured andoperated in substantially the same manner as bubble column reactor 20described above with reference to FIG. 1.

In one or more embodiments, primary oxidation reactor 102 and secondaryoxidation reactor 104 are bubble column reactors. Primary oxidationreactor 102 can include a primary reaction vessel 106 and a primaryoxidant sparger 108, while secondary oxidation reactor 104 can include asecondary reaction vessel 110 and a lower oxidant sparger 112. Asdiscussed in greater detail below, secondary oxidation reactor 104 canoptionally comprise one or more upper oxidant spargers as well. In oneor more embodiments, primary and secondary reaction vessels 106 and 110can each include a respective upright sidewall having a generallycylindrical configuration. The ratio of the maximum height of theupright sidewall of secondary reaction vessel 110 to the maximum heightof the upright sidewall of primary reaction vessel 106 can be in therange of from about 0.1:1 to about 0.9:1, in the range of from about0.2:1 to about 0.8:1, or in the range of from 0.3:1 to 0.7:1.

Primary reaction vessel 106 defines therein a primary reaction zone 116,while secondary reaction vessel 110 defines therein a secondary reactionzone 118. In various embodiments, the ratio of the maximum horizontalcross sectional area of secondary reaction zone 118 to primary reactionzone 116 can be in the range of from about 0.01:1 to about 0.75:1, inthe range of from about 0.02:1 to about 0.5:1, or in the range of from0.04:1 to 0.3:1. Additionally, the volume ratio of primary reaction zone116 to secondary reaction zone 118 can be in the range of from about 1:1to about 100:1, in the range of from about 4:1 to about 50:1, or in therange of from 8:1 to 30:1. Furthermore, primary reaction zone 116 canhave a ratio of maximum vertical height to maximum horizontal diameterin the range of from about 3:1 to about 30:1, in the range of from about6:1 to about 20:1, or in the range of from 9:1 to 15:1.

As shown in FIG. 2, secondary reaction zone 118 can have a maximumvertical length L_(s) and a maximum horizontal diameter D_(s). In one ormore embodiments, secondary reaction zone 118 can have a ratio ofmaximum vertical length to maximum horizontal diameter “L_(s):D_(s)” inthe range of from about 14:1 to about 28:1, in the range of from about16:1 to about 26:1, in the range of from about 18:1 to about 24:1, inthe range of from about 20:1 to about 23:1, or in the range of from 21:1to 22:1. In various embodiments, D_(s) of secondary reaction zone 118can be in the range of from about 0.1 to about 5 meters, in the range offrom about 0.3 to about 4 meters, or in the range of from 1 to 3 meters.Furthermore, L_(s) of secondary reaction zone 118 can be in the range offrom about 1 to about 100 meters, in the range of from about 3 to about50 meters, or in the range of from 10 to 40 meters.

As with bubble column reactor 20 described above with respect to FIG. 1,primary reaction zone 116 has a maximum vertical length L_(p) and amaximum horizontal diameter D_(p). In various embodiments, the ratio ofthe maximum horizontal diameter of secondary reaction zone 118 to themaximum horizontal diameter of primary reaction zone 116 “D_(s):D_(p)”can be in the range of from about 0.05:1 to about 0.8:1, in the range offrom about 0.1:1 to about 0.6:1, or in the range of from 0.2:1 to 0.5:1.Furthermore, the ratio of the maximum vertical length of secondaryreaction zone 118 to the maximum vertical length of primary reactionzone 116 “L_(s):L_(p)” can be in the range of from about 0.03:1 to about1:1, in the range of from about 0.1:1 to about 0.9:1, or in the range offrom 0.3:1 to 0.8:1.

In various embodiments, secondary oxidation reactor 104 can be locatedalongside primary oxidation reactor 102 (i.e., at least a portion ofprimary and secondary oxidation reactors 102 and 104 share a commonelevation). As noted above, primary reaction zone 116 of primaryoxidation reactor 102 has a maximum diameter D_(p). In one or moreembodiments, the volumetric centroid of secondary reaction zone 118 canbe horizontally spaced from the volumetric centroid of primary reactionzone 416 by at least about 0.5D_(p), 0.75D_(p), or 1.0D_(p) and by lessthan about 30D_(p), 10D_(p), or 3D_(p).

Any parameters (e.g., height, width, area, volume, relative horizontalplacement, and relative vertical placement) specified herein for primaryreaction vessel 106 and appurtenances are also construed as applying toprimary reaction zone 116 defined by primary reaction vessel 106, andvice versa. Further, any parameters specified herein for secondaryreaction vessel 110 and appurtenances are also construed as applying tosecondary reaction zone 118 defined by secondary reaction vessel 110,and vice versa.

During normal operation of reactor system 100, reaction medium 120 canfirst undergo oxidation in primary reaction zone 116 of primaryoxidation reactor 102. Reaction medium 120 a can then be withdrawn fromprimary reaction zone 116 and transferred to secondary reaction zone 118via conduit 105. In secondary reaction zone 118, the liquid and/or solidphases of reaction medium 120 b can be subjected to further oxidation.In various embodiments, at least about 50, 75, 95, or 99 weight percentof liquid and/or solid phases withdrawn from primary reaction zone 116can be processed in secondary reaction zone 116. Overhead gasses canexit an upper gas outlet of secondary oxidation reactor 104 and can betransferred back to primary oxidation reactor 102 via conduit 107. Aslurry phase of reaction medium 120 b can exit a lower slurry outlet 122of secondary oxidation reactor 104 and can thereafter be subjected tofurther downstream processing.

Inlet conduit 105 may attach to primary oxidation reactor 102 at anyheight. Although not shown in FIG. 2, reaction medium 120 can bemechanically pumped to secondary reaction zone 118 if desired. However,elevation head (gravity) can also be used transfer reaction medium 120from primary reaction zone 116 through inlet conduit 105 and intosecondary reaction zone 118. Accordingly, inlet conduit 105 can beconnected on one end to the upper 50, 30, 20, or 10 percent of the totalheight and/or volume of primary reaction zone 116. In other variousembodiments, the slurry outlet (not depicted) through which reactionmedium 120 a can exit primary oxidation reactor 102 into inlet conduit105 can be spaced at least 0.1L_(p), at least 0.2L_(p), or at least0.3L_(p) away from each of the normally top and normally bottom ends ofprimary reaction zone 116.

In various embodiments, the other end of inlet conduit 105 can beattached in fluid flow communication to a slurry inlet (not depicted)located in the upper 30, 20, 10, or 5 percent of the total height and/orvolume of secondary reaction zone 118. In alternate embodiments, theslurry inlet in secondary oxidation reactor 104 can be a mid-levelslurry inlet spaced from the bottom of secondary reaction zone 118 by adistance in the range of from about 0.3L_(s) to about 0.9L_(s), in therange of from about 0.4L_(s) to about 0.8L_(s), in the range of fromabout 0.5L_(s) to about 0.8L_(s), or in the range of from 0.55L_(s) to0.6L_(s). Additionally, the slurry inlet in secondary oxidation reactor104 can be spaced from the bottom of the secondary reaction zone by adistance in the range of from about 9D_(s) to about 15D_(s), in therange of from about 10D, to about 14D_(s), or in the range of from11D_(s) to 13D_(s). In operation, at least a portion of reaction medium120 a can be introduced into secondary reaction zone 118 via themid-level slurry inlet. In various embodiments, at least 5 volumepercent, at least 10 volume percent, at least 20 volume percent, atleast 30 volume percent, at least 50 volume percent, at least 75 volumepercent, or 100 volume percent of the total amount of reaction medium120 a introduced into secondary reaction zone 118 can be introduced viathe mid-level slurry inlet.

In various embodiments, inlet conduit 105 can be horizontal,substantially horizontal, and/or sloping downward from primary oxidationreactor 102 toward secondary oxidation reactor 104. In one or moreembodiments, inlet conduit 105 is horizontal or substantiallyhorizontal, and can be straight or substantially straight. Accordingly,in one or more embodiments, the slurry outlet (not depicted) from theprimary oxidation reactor 102 can be at the same or substantially thesame vertical elevation as the slurry inlet (not depicted) in secondaryoxidation reactor 104.

In various embodiments, outlet conduit 107 may attach to any elevationin secondary oxidation reactor 104. In various embodiments, outletconduit 107 can be connected to secondary oxidation reactor 104 abovethe attachment elevation of inlet conduit 105. Furthermore, outletconduit 107 can attach to the top of secondary oxidation reactor 104.Outlet conduit 107 can attach to primary oxidation reactor 102 above theattachment elevation of inlet conduit 105. In various embodiments,outlet conduit 107 attaches to the upper 30, 20, 10, or 5 percent of thetotal height and/or volume of primary reaction zone 116. Outlet conduit107 can be horizontal and/or sloping upward from secondary oxidationreactor 104 toward primary oxidation reactor 102. Although not shown inFIG. 2, outlet conduit 107 may also attach directly to the gas outletconduit that withdraws gaseous effluent from the top of primaryoxidation reactor 102.

The upper extent of secondary reaction zone 116 may be above or belowthe upper extent of primary reaction zone 118. In various embodiments,the upper extent of primary reaction zone 116 can be within 10 metersabove to 50 meters below, 2 meters below to 40 meters below, or 5 metersbelow to 30 meters below the upper extent of secondary reaction zone118. The lower extent of secondary reaction zone 118 may be elevatedabove or below the lower extent of primary reaction zone 116. In variousembodiments, the lower extent of primary reaction zone 116 can beelevated within about 40, 20, 5, or 2 meters above or below the lowerextent of secondary reaction zone 118.

Lower slurry outlet 122 may exit from any elevation of secondaryoxidation reactor 104. In various embodiments, lower slurry outlet 122can be connected to secondary oxidation reactor 104 below the attachmentelevation of inlet conduit 105. In various embodiments, lower slurryoutlet 122 attaches to the bottom of secondary oxidation reactor 104 asshown in FIG. 2.

Secondary oxidation reactor 104 can comprise at least one oxidant inletthat permits additional molecular oxygen to be discharged into secondaryreaction zone 118. In one or more embodiments, secondary oxidationreactor 104 can comprise at least one normally lower oxidant inlet andat least one normally upper oxidant inlet. In various embodiments, thenormally lower oxidant inlet can be spaced from the bottom of secondaryreaction zone 118 by less than 0.5L_(s), less than 0.4L_(s), less than0.3L_(s), or less than 0.24. Additionally, the normally upper oxidantinlet can be spaced from the bottom of secondary reaction zone 118 by atleast 0.5L_(s), at least 0.6L_(s), at least 0.7L_(s), at least 0.8L_(s),or at least 0.94. In one or more embodiments, secondary oxidationreactor 104 can comprise at least two normally upper oxidant inlets,each spaced from the bottom of the secondary reaction zone 118 by atleast 0.5L_(s), at least 0.55L_(s), at least 0.6L_(s), at least0.7L_(s), at least 0.8L_(s), or at least 0.94. Additionally, as notedabove, secondary oxidation reactor 104 can comprise a slurry inlet thatis in fluid-flow communication with inlet conduit 105. In variousembodiments, the normally upper oxidant inlet can be spaced less than0.4L_(s), less than 0.3L_(s), less than 0.2L_(s), or less than 0.14 fromthe slurry inlet in secondary oxidation reactor 104. In otherembodiments, the normally upper oxidant inlet can be spaced above theslurry inlet by less than 0.4L_(s), less than 0.3L_(s), less than0.2L_(s), or less than 0.1L_(s).

During operation, a first portion of the gas-phase oxidant introducedinto secondary reaction zone 118 can be introduced via the normallyupper oxidant inlet, while a second portion of the gas-phase oxidant canbe introduced via the normally lower oxidant inlet. In variousembodiments, the first portion of the gas-phase oxidant introduced viathe normally upper oxidant inlet can constitute in the range of fromabout 5 to about 49 percent, in the range of from about 5 to about 35percent, in the range of from about 10 to about 20 percent, or in therange of from 10 to 15 percent of the total volume of gas-phase oxidantintroduced into secondary reaction zone 118. Accordingly, the normallyupper oxidant inlet and normally lower oxidant inlet can define betweenthem a total open area for introducing gas-phase oxidant into secondaryreaction zone 118. In one or more embodiments, the normally upperoxidant inlet can define in the range of from about 5 to about 49percent of the total open area, in the range of from about 5 to about 35percent of the total open area, in the range of from about 10 to about20 percent of the total open area, or in the range of from 10 to 15percent of the total open area.

As shown in FIG. 2, the above-mentioned lower oxidant inlet can comprisea lower oxidant sparger 112. Additionally the above-mentioned upperoxidant inlet(s) can comprise one or more upper oxidant spargers 114a,b,c. Referring now to FIG. 3, a cross-section of secondary oxidationreactor 104 is shown along line 3-3, particularly illustrating upperoxidant sparger 114 a. As seen in FIG. 3, upper oxidant sparger 114 acan comprise a plurality of oxidant discharge openings 124 forintroducing gas-phase oxidant into secondary reaction zone 118. Althoughnot shown, each of upper oxidant spargers 114 b and 114 c can alsocomprise a plurality of oxidant discharge openings. Similarly, loweroxidant sparger 112 can also comprise a plurality of oxidant dischargeopenings. In one or more embodiments, at least 50 percent, at least 60percent, at least 70 percent, at least 80 percent, at least 90 percent,at least 95 percent or at least 99 percent of oxidant discharge openings124 defined by upper oxidant spargers 114 a,b,c can be oriented todischarge a gas-phase oxidant in the normally downward direction. Asused herein, the term “downward” shall denote any direction extendingbelow the normally underneath side of upper oxidant spargers 114 a,b,cwithin 15° of vertical. In various embodiments, at least 50 percent, atleast 60 percent, at least 70 percent, at least 80 percent, at least 90percent, at least 95 percent, or at least 99 percent of oxidantdischarge openings located in lower oxidant sparger 112 can be orientedto discharge gas-phase oxidant in a normally downward direction and/orat a 45° angle or approximately a 45° angle away from verticallydownward.

As noted above, at least a portion of the gas-phase oxidant and thereaction medium 120 a introduced into secondary reaction zone 118 cancombine to form reaction medium 120 b. In one or more embodiments, itmay be desirable for reaction medium 120 b to have minimal zones of lowoxygen concentration. This minimization of low oxygen content zones canbe quantified by theoretically partitioning the entire volume ofreaction medium 120 b into 20 discrete horizontal slices of uniformvolume. With the exception of the highest and lowest horizontal slices,each horizontal slice is a discrete volume bounded on its sides by thesidewall of the reactor and bounded on its top and bottom by imaginaryhorizontal planes. The highest horizontal slice is bounded on its bottomby an imaginary horizontal plane and on its top by the upper surface ofthe reaction medium or, in the case of a liquid-full column, by theupper end of the vessel. The lowest horizontal slice is bounded on itstop by an imaginary horizontal plane and on its bottom by the lower endof the vessel. In various embodiments, when the entire volume ofreaction medium 120 b is theoretically partitioned into 20 discretehorizontal slices of equal volume, no two adjacent horizontal sliceshave a combined time-averaged and volume-averaged oxygen content of lessthan 7, less than 8, less than 9, or less than 10 ppmw. In otherembodiments, none of the 20 horizontal slices has a time-averaged andvolume-averaged oxygen content of less than 7, less than 8, less than 9,or less than 10 ppmw.

Referring again to FIG. 2, in general, the manner in which the feed,oxidant, and reflux streams are introduced into primary oxidationreactor 102 and the manner in which primary oxidation reactor 102 isoperated are substantially the same as described above with reference tobubble column reactor 20 of FIG. 1. However, one difference betweenprimary oxidation reactor 102 (FIG. 2) and bubble column reactor 20(FIG. 1) is that primary oxidation reactor 102 does not include anoutlet that permits the slurry phase of reaction medium 120 a to bedirectly discharged from primary reaction vessel 106 for downstreamprocessing. Rather, primary oxidation reactor 102 requires the slurryphase of reaction medium 120 a to first pass through secondary oxidationreactor 104 before being discharged from reactor system 100. Asmentioned above, in secondary reaction zone 118 of secondary oxidationreactor 104, reaction medium 120 b is subjected to further oxidation tohelp purify the liquid and/or solid phases of reaction medium 120 b.

In a process where para-xylene is fed to reaction zone 116, the liquidphase of reaction medium 120 a that exits primary reaction zone 116 andenters secondary reaction zone 118 typically contains at least somepara-toluic acid. In various embodiments, a substantial portion of thepara-toluic acid entering secondary reaction zone 118 can be oxidized insecondary reaction zone 118. Thus, the time-averaged concentration ofpara-toluic acid in the liquid phase of reaction medium 120 b exitingsecond reaction zone 118 can be less than the time-averagedconcentration of para-toluic acid in the liquid phase of reaction medium120 a/b entering secondary reaction zone 118. In various embodiments,the time-averaged concentration of para-toluic acid in the liquid phaseof reaction medium 120 b exiting secondary reaction zone 118 can be lessthan about 50, 10, or 5 percent of the time-averaged concentration ofpara-toluic acid in the liquid phase of reaction medium 120 a/b enteringsecondary reaction zone 118. The time-averaged concentration ofpara-toluic acid in the liquid phase of reaction medium 120 a/b enteringsecond reaction zone 118 can be at least about 250 ppmw, in the range offrom about 500 to about 6,000 ppmw, or in the range of from 1,000 to4,000 ppmw. By comparison, the time-averaged concentration ofpara-toluic acid in the liquid phase of reaction medium 120 b exitingsecondary reaction zone 118 can be less than about 1,000, 250, or 50ppmw.

As reaction medium 120 b is processed in secondary reaction zone 118 ofsecondary oxidation reactor 104, the gas hold-up of reaction medium 120b can decrease as the slurry phase of reaction medium 120 b flowsdownwardly through secondary reaction zone 118. In various embodiments,the ratio of the time-averaged gas hold-up of reaction medium 120 a/bentering secondary reaction zone 118 to reaction medium 120 b exitingsecondary reaction zone 118 can be at least about 2:1, 10:1, or 25:1.Additionally, the time-averaged gas hold-up of reaction medium 120 a/bentering secondary reaction zone 118 can be in the range of from about0.4 to about 0.9, in the range of from about 0.5 to about 0.8, or in therange of from 0.55 to 0.7. Furthermore, the time-averaged gas hold-up ofreaction medium 120 b exiting secondary reaction zone 118 can be lessthan about 0.1, 0.05, or 0.02. In one or more embodiments, the ratio ofthe time-averaged gas hold-up of reaction medium 120 a in primaryreaction zone 116 to reaction medium 120 b in secondary reaction zone118 can be greater than about 1:1, in the range of from about 1.25:1 toabout 5:1, or in the range of from 1.5:1 to 4:1, where the gas hold-upvalues are measured at any height of primary and secondary reactionzones 116 and 118, at any corresponding heights of primary and secondaryreaction zones 116 and 118, at ¼-height of primary and/or secondaryreaction zones 116 and 118, at ½-height of primary and/or secondaryreaction zones 116 and 118, at ¾-height of primary and/or secondaryreaction zones 116 and 118, and/or are average values over the entireheights of primary and/or secondary reaction zones 116 and 118. Invarious embodiments, the time-averaged gas hold-up of the portion ofreaction medium 120 a in primary reaction zone 116 can be in the rangeof from about 0.4 to about 0.9, in the range of from about 0.5 to about0.8, or in the range of from 0.55 to 0.70, where the gas hold-up ismeasured at any height of primary reaction zone 116, at ¼-height ofprimary reaction zone 116, at ½-height of primary reaction zone 116, at¾-height of primary reaction zone 116, and/or is an average over theentire height of primary reaction zone 116. Additionally, thetime-averaged gas hold-up of the portion of reaction medium 120 b insecondary reaction zone 118 can be in the range of from about 0.01 toabout 0.6, in the range of from about 0.03 to about 0.3, or in the rangeof from 0.08 to 0.2, where the gas hold-up is measured at any height ofsecondary reaction zone 118, at ¼-height of secondary reaction zone 118,at ½-height of secondary reaction zone 118, at ¾-height of secondaryreaction zone 118, and/or is an average over the entire height ofsecondary reaction zone 118.

The temperature of reaction medium 120 can be approximately the same inprimary and secondary reaction zones 116 and 118. In variousembodiments, such temperature can be in the range of from about 125 toabout 200° C., in the range of from about 140 to about 180° C., or inthe range of from 150 to 170° C. However, temperature differences can beformed within primary reaction zone 116, such as those described ingreater detail below with reference to FIG. 4. In various embodiments,temperature differences of the same magnitudes can also exist withinsecondary reaction zone 118 and also between primary reaction zone 116and secondary reaction zone 118. These additional temperature gradientsrelate to chemical reaction occurring in secondary reaction zone 118,the introduction of additional oxidant to secondary reaction zone 118,and the static pressures extant in secondary reaction zone 118 comparedto those in primary reaction zone 116. As disclosed above, in variousembodiments, the bubble hold-up can be greater in primary reaction zone116 than in secondary reaction zone 118. Thus, the static pressure inprimary reaction zone 116 can be greater than in secondary reaction zone118. The magnitude of this pressure difference depends on the magnitudeof liquid or slurry density and on the difference in bubble hold-upbetween the two reaction zones. The magnitude of this pressuredifference increases at elevations further below the upper boundary ofsecondary reaction zone 118.

As seen in FIG. 2, a portion of the total molecular oxygen fed toreactor system 100 is introduced into secondary reaction zone 118 ofsecondary oxidation reactor 104 via lower oxidant sparger 112 andoptionally via one or more of upper oxidant spargers 114 a,b,c. Invarious embodiments, the majority of the total molecular oxygen fed toreactor system 100 can be introduced into primary reaction zone 116,with the balance being introduced into secondary reaction zone 118. Inone or more embodiments, at least about 70, 90, 95, or 98 mole percentof the total molecular oxygen fed to reactor system 100 can beintroduced into primary reaction zone 116. Furthermore, the molar ratioof the amount of molecular oxygen introduced into primary reaction zone116 to the amount of molecular oxygen introduced into secondary reactionzone 118 can be at least about 2:1, in the range of from about 4:1 toabout 200:1, or in the range of from 10:1 to 100:1. Although it ispossible for some of the solvent and/or oxidizable compound (e.g.,para-xylene) to be fed directly to secondary reaction zone 118, invarious embodiments, less than about 10, 5, or 1 weight percent of thetotal amount of solvent and/or oxidizable compound fed to reactor system100 is fed directly to secondary reaction zone 118.

The volume, residence time, and space time rate of reaction medium 120 ain primary reaction zone 116 of primary reaction vessel 106 can be, invarious embodiments, substantially greater than the volume, residencetime, and space time rate of reaction medium 120 b in secondary reactionzone 118 of secondary reaction vessel 110. Therefore, the majority ofthe oxidizable compound (e.g., para-xylene) fed to reactor system 100can be oxidized in primary reaction zone 116. In various embodiments, atleast about 80, 90, or 95 weight percent of all the oxidizable compoundthat is oxidized in reactor system 100 can be oxidized in primaryreaction zone 116.

In one or more embodiments, the time-averaged superficial gas velocityof reaction medium 120 a in primary reaction zone 116 can be at leastabout 0.2, 0.4, 0.8, or 1 meter per second, where the superficial gasvelocity is measured at any height of primary reaction zone 116, at¼-height of primary reaction zone 116, at ½-height of primary reactionzone 116, at ¾-height of primary reaction zone 116, and/or is an averageover the entire height of primary reaction zone 116. Although reactionmedium 120 b in secondary reaction zone 118 can have the samesuperficial gas velocity as reaction medium 120 a in primary reactionzone 116, in various embodiments the time-averaged superficial gasvelocity of reaction medium 120 b in secondary reaction zone 118 can beless than the time-averaged superficial gas velocity of reaction medium120 a in primary reaction zone 116. This reduced superficial gasvelocity in secondary reaction zone 118 is made possible by, forexample, the reduced demand for molecular oxygen in secondary reactionzone 118 compared to primary reaction zone 116. The ratio of thetime-averaged superficial gas velocity of reaction medium 120 a inprimary reaction zone 116 to reaction medium 120 b in secondary reactionzone 118 can be at least about 1.25:1, 2:1, or 5:1, where thesuperficial gas velocities are measured at any height of primary andsecondary reaction zones 116 and 118, at any corresponding heights ofprimary and secondary reaction zones 116 and 118, at ¼-height of primaryand/or secondary reaction zones 116 and 118, at ½-height of primaryand/or secondary reaction zones 116 and 118, at ¾-height of primaryand/or secondary reaction zones 116 and 118, and/or are average valuesover the entire heights of primary and/or secondary reaction zones 116and 118. In various embodiments, the time-averaged and volume-averagedsuperficial gas velocity of reaction medium 120 b in secondary reactionzone 118 can be less than about 0.2, 0.1, or 0.06 meters per second,where the superficial gas velocity is measured at any height ofsecondary reaction zone 118, at ¼-height of secondary reaction zone 118,at ½-height of secondary reaction zone 118, at ¾-height of secondaryreaction zone 118, and/or is an average over the entire height ofsecondary reaction zone 118. With these lower superficial gasvelocities, downward flow of the slurry phase of reaction medium 120 bin secondary reaction zone 118 can be made to move directionally towardplug flow. For example, during oxidation of para-xylene to form TPA, therelative vertical gradient of liquid phase concentration of para-toluicacid can be much greater in secondary reaction zone 118 than in primaryreaction zone 116. This is notwithstanding that secondary reaction zone118 is a bubble column having axial mixing of liquid and of slurrycompositions. The time-averaged superficial velocity of the slurry phase(solid+liquid) and the liquid phase of reaction medium 120 b insecondary reaction zone 118 can be less than about 0.2, 0.1, or 0.06meters per second, where the superficial velocity is measured at anyheight of secondary reaction zone 118, at ¼-height of secondary reactionzone 118, at ½-height of secondary reaction zone 118, at ¾-height ofsecondary reaction zone 118, and/or is an average over the entire heightof secondary reaction zone 118.

In various embodiments, the liquid phase of reaction medium 120 blocated in secondary reaction zone 118 can have a mass-averagedresidence time in secondary reaction zone 118 of at least about 1minute, in the range of from about 2 to about 60 minutes, or in therange of from 5 to 30 minutes.

As mentioned above, certain physical and operational features of thebubble column reactors, described above with reference to FIG. 1,provide for vertical gradients in the pressure, temperature, andreactant (i.e., oxygen and oxidizable compound) concentrations of theprocessed reaction medium. As discussed above, these vertical gradientscan provide for a more effective and economical oxidation process ascompared to conventional oxidation processes, which favor a well-mixedreaction medium of relatively uniform pressure, temperature, andreactant concentration throughout. The vertical gradients for oxygen,oxidizable compound (e.g., para-xylene), and temperature made possibleby employing an oxidation system in accordance with an embodiment of thepresent invention will now be discussed in greater detail.

Referring now to FIG. 4, in order to quantify the reactant concentrationgradients existing in the reaction medium during oxidation in the bubblecolumn reactor, the entire volume of the reaction medium can betheoretically partitioned into 30 discrete horizontal slices of equalvolume. FIG. 4 illustrates the concept of dividing the reaction mediuminto 30 discrete horizontal slices of equal volume. With the exceptionof the highest and lowest horizontal slices, each horizontal slice is adiscrete volume bounded on its top and bottom by imaginary horizontalplanes and bounded on its sides by the wall of the reactor. The highesthorizontal slice is bounded on its bottom by an imaginary horizontalplane and on its top by the upper surface of the reaction medium. Thelowest horizontal slice is bounded on its top by an imaginary horizontalplane and on its bottom by the bottom of the vessel shell. Once thereaction medium has been theoretically partitioned into 30 discretehorizontal slices of equal volume, the time-averaged and volume-averagedconcentration of each horizontal slice can then be determined. Theindividual horizontal slice having the maximum concentration of all 30horizontal slices can be identified as the “C-max horizontal slice.” Theindividual horizontal slice located above the C-max horizontal slice andhaving the minimum concentration of all horizontal slices located abovethe C-max horizontal slice can be identified as the “C-min horizontalslice.” The vertical concentration gradient can then be calculated asthe ratio of the concentration in the C-max horizontal slice to theconcentration in the C-min horizontal slice.

With respect to quantifying the oxygen concentration gradient, when thereaction medium is theoretically partitioned into 30 discrete horizontalslices of equal volume, an O₂-max horizontal slice is identified ashaving the maximum oxygen concentration of all the 30 horizontal slicesand an O₂-min horizontal slice is identified as having the minimumoxygen concentration of the horizontal slices located above the O₂-maxhorizontal slice. The oxygen concentrations of the horizontal slices aremeasured in the gas phase of the reaction medium on a time-averaged andvolume-averaged molar wet basis. In various embodiments, the ratio ofthe oxygen concentration of the O₂-max horizontal slice to the oxygenconcentration of the O₂-min horizontal slice can be in the range of fromabout 2:1 to about 25:1, in the range of from about 3:1 to about 15:1,or in the range of from 4:1 to 10:1.

Typically, the O₂-max horizontal slice will be located near the bottomof the reaction medium, while the O₂-min horizontal slice will belocated near the top of the reaction medium. In one or more embodiments,the O₂-min horizontal slice can be one of the 5 upper-most horizontalslices of the 30 discrete horizontal slices. Additionally, the O₂-minhorizontal slice can be the upper-most one of the 30 discrete horizontalslices, as illustrated in FIG. 4. In various embodiments, the O₂-maxhorizontal slice can be one of the 10 lower-most horizontal slices ofthe 30 discrete horizontal slices. Additionally, the O₂-max horizontalslice can be one of the 5 lower-most horizontal slices of the 30discrete horizontal slices. For example, FIG. 4 illustrates the O₂-maxhorizontal slice as the third horizontal slice from the bottom of thereactor. In one or more embodiments, the vertical spacing between theO₂-min and O₂-max horizontal slices can be at least about 2W_(p), atleast about 4W_(p), or at least 6W_(p). Additionally, the verticalspacing between the O₂-min and O₂-max horizontal slices can be at leastabout 0.2H_(p), at least about 0.4H_(p), or at least 0.6H_(p).

The time-averaged and volume-averaged oxygen concentration, on a wetbasis, of the O₂-min horizontal slice can be in the range of from about0.1 to about 3 mole percent, in the range of from about 0.3 to about 2mole percent, or in the range of from 0.5 to 1.5 mole percent. Thetime-averaged and volume-averaged oxygen concentration of the O₂-maxhorizontal slice can be in the range of from about 4 to about 20 molepercent, in the range of from about 5 to about 15 mole percent, or inthe range of from 6 to 12 mole percent. The time-averaged concentrationof oxygen, on a dry basis, in the gaseous effluent discharged from thereactor via the gas outlet can be in the range of from about 0.5 toabout 9 mole percent, in the range of from about 1 to about 7 molepercent, or in the range of from 1.5 to 5 mole percent.

Because the oxygen concentration decays so markedly toward the top ofthe reaction medium, the demand for oxygen can be reduced in the top ofthe reaction medium. This reduced demand for oxygen near the top of thereaction medium can be accomplished by creating a vertical gradient inthe concentration of the oxidizable compound (e.g., para-xylene), wherethe minimum concentration of oxidizable compound is located near the topof the reaction medium.

With respect to quantifying the oxidizable compound (e.g., para-xylene)concentration gradient, when the reaction medium is theoreticallypartitioned into 30 discrete horizontal slices of equal volume, anOC-max horizontal slice is identified as having the maximum oxidizablecompound concentration of all the 30 horizontal slices and an OC-minhorizontal slice is identified as having the minimum oxidizable compoundconcentration of the horizontal slices located above the OC-maxhorizontal slice. The oxidizable compound concentrations of thehorizontal slices are measured in the liquid phase on a time-averagedand volume-averaged mass fraction basis. In various embodiments, theratio of the oxidizable compound concentration of the OC-max horizontalslice to the oxidizable compound concentration of the OC-min horizontalslice can be greater than about 5:1, greater than about 10:1, greaterthan about 20:1, or in the range of from 40:1 to 1000:1.

Typically, the OC-max horizontal slice will be located near the bottomof the reaction medium, while the OC-min horizontal slice will belocated near the top of the reaction medium. In one or more embodiments,the OC-min horizontal slice can be one of the 5 upper-most horizontalslices of the 30 discrete horizontal slices. Additionally, the OC-minhorizontal slice can be the upper-most one of the 30 discrete horizontalslices, as illustrated in FIG. 4. In various embodiments, the OC-maxhorizontal slice can be one of the lower-most horizontal slices of the30 discrete horizontal slices. Additionally, the OC-max horizontal slicecan be one of the 5 lower-most horizontal slices of the 30 discretehorizontal slices. For example, FIG. 4 illustrates the OC-max horizontalslice as the fifth horizontal slice from the bottom of the reactor. Invarious embodiments, the vertical spacing between the OC-min and OC-maxhorizontal slices can be at least about 2W_(p) (where “W_(p)” is themaximum width of the reaction medium), at least about 4W_(p), or atleast 6W_(p). Given a height “H_(p)” of the reaction medium, thevertical spacing between the OC-min and OC-max horizontal slices can beat least about 0.2H_(p), at least about 0.4H_(p), or at least 0.6H_(p).

The time-averaged and volume-averaged oxidizable compound (e.g.,para-xylene) concentration in the liquid phase of the OC-min horizontalslice can be less than about 5,000 ppmw, less than about 2,000 ppmw,less than about 400 ppmw, or in the range of from 1 ppmw to 100 ppmw.The time-averaged and volume-averaged oxidizable compound concentrationin the liquid phase of the OC-max horizontal slice can be in the rangeof from about 100 ppmw to about 10,000 ppmw, in the range of from about200 ppmw to about 5,000 ppmw, or in the range of from 500 ppmw to 3,000ppmw.

Although the bubble column reactor can provide vertical gradients in theconcentration of the oxidizable compound, the volume percent of thereaction medium having an oxidizable compound concentration in theliquid phase above 1,000 ppmw can also be minimized. In variousembodiments, the time-averaged volume percent of the reaction mediumhaving an oxidizable compound concentration in the liquid phase above1,000 ppmw can be less than about 9 percent, less than about 6 percent,or less than 3 percent. Additionally, the time-averaged volume percentof the reaction medium having an oxidizable compound concentration inthe liquid phase above 2,500 ppmw can be less than about 1.5 percent,less than about 1 percent, or less than 0.5 percent. Furthermore, thetime-averaged volume percent of the reaction medium having an oxidizablecompound concentration in the liquid phase above 10,000 ppmw can be lessthan about 0.3 percent, less than about 0.1 percent, or less than 0.03percent. Also, the time-averaged volume percent of the reaction mediumhaving an oxidizable compound concentration in the liquid phase above25,000 ppmw can be less than about 0.03 percent, less than about 0.015percent, or less than 0.007 percent. The inventors note that the volumeof the reaction medium having the elevated levels of oxidizable compoundneed not lie in a single contiguous volume. At many times, the chaoticflow patterns in a bubble column reaction vessel produce simultaneouslytwo or more continuous but segregated portions of the reaction mediumhaving the elevated levels of oxidizable compound. At each time used inthe time averaging, all such continuous but segregated volumes largerthan 0.0001 volume percent of the total reaction medium are addedtogether to determine the total volume having the elevated levels ofoxidizable compound concentration in the liquid phase.

In addition to the concentration gradients of oxygen and oxidizablecompound, discussed above, a temperature gradient can exist in thereaction medium. Referring again to FIG. 4, this temperature gradientcan be quantified in a manner similar to the concentration gradients bytheoretically partitioning the reaction medium into 30 discretehorizontal slices of equal volume and measuring the time-averaged andvolume-averaged temperature of each slice. The horizontal slice with thelowest temperature out of the lowest 15 horizontal slices can then beidentified as the T-min horizontal slice, and the horizontal slicelocated above the T-min horizontal slice and having the maximumtemperature of all the slices above the T-min horizontal slice can thenbe identified as the T-max horizontal slice. In various embodiments, thetemperature of the T-max horizontal slice can be at least about 1° C.higher than the temperature of the T-min horizontal slice, in the rangeof from about 1.25 to about 12° C. higher than the temperature of theT-min horizontal slice, or in the range of from 2 to 8° C. higher thanthe temperature of the T-min horizontal slice. The temperature of theT-max horizontal slice can be in the range of from about 125 to about200° C., in the range of from about 140 to about 180° C., or in therange of from 150 to 170° C.

Typically, the T-max horizontal slice will be located near the center ofthe reaction medium, while the T-min horizontal slice will be locatednear the bottom of the reaction medium. In various embodiments, theT-min horizontal slice can be one of the 10 lower-most horizontal slicesof the 15 lowest horizontal slices, or one of the 5 lower-mosthorizontal slices of the 15 lowest horizontal slices. For example, FIG.4 illustrates the T-min horizontal slice as the second horizontal slicefrom the bottom of the reactor. In various embodiments, the T-maxhorizontal slice can be one of the 20 middle horizontal slices of thediscrete horizontal slices, or one of the 14 middle horizontal slices ofthe 30 discrete horizontal slices. For example, FIG. 4 illustrates theT-max horizontal slice as the twentieth horizontal slice from the bottomof the reactor (i.e., one of the middle 10 horizontal slices). Thevertical spacing between the T-min and T-max horizontal slices can be atleast about 2W_(p), at least about 4W_(p), or at least 6W_(p). Thevertical spacing between the T-min and T-max horizontal slices can be atleast about 0.2H_(p), at least about 0.4H_(p), or at least 0.6H_(p).

As discussed above, when a vertical temperature gradient exists in thereaction medium, it can be advantageous to withdraw the reaction mediumat an elevated location where the temperature of reaction medium ishighest, especially when the withdrawn product is subjected to furtherdownstream processing at higher temperatures. Thus, when reaction medium120 is withdrawn from the reaction zone via one or more elevatedoutlets, as illustrated in FIG. 2, the elevated outlet(s) can be locatednear the T-max horizontal slice. In various embodiments, the elevatedoutlet can be located within 10 horizontal slices of the T-maxhorizontal slice, within 5 horizontal slices of the T-max horizontalslice, or within 2 horizontal slices of the T-max horizontal slice.

It is now noted that many of the inventive features described herein canbe employed in multiple oxidation reactor systems—not just systemsemploying a single oxidation reactor. In addition, certain inventivefeatures described herein can be employed in mechanically-agitatedand/or flow-agitated oxidation reactors—not just bubble-agitatedreactors (i.e., bubble column reactors). For example, the inventors havediscovered certain advantages associated with staging/varying oxygenconcentration and/or oxygen consumption rate throughout the reactionmedium. The advantages realized by the staging of oxygenconcentration/consumption in the reaction medium can be realized whetherthe total volume of the reaction medium is contained in a single vesselor in multiple vessels. Further, the advantages realized by the stagingof oxygen concentration/consumption in the reaction medium can berealized whether the reaction vessel(s) is mechanically-agitated,flow-agitated, and/or bubble-agitated.

One way of quantifying the degree of staging of oxygen concentrationand/or consumption rate in a reaction medium is to compare two or moredistinct 20-percent continuous volumes of the reaction medium. These20-percent continuous volumes need not be defined by any particularshape. However, each 20-percent continuous volume must be formed of acontiguous volume of the reaction medium (i.e., each volume is“continuous”), and the 20-percent continuous volumes must not overlapone another (i.e., the volumes are “distinct”). These distinct20-percent continuous volumes can be located in the same reactor or inmultiple reactors. Referring now to FIG. 5, the bubble column reactor isillustrated as containing a reaction medium that includes a firstdistinct 20-percent continuous volume 37 and a second distinct20-percent continuous volume 39.

The staging of oxygen availability in the reaction medium can bequantified by referring to the 20-percent continuous volume of reactionmedium having the most abundant mole fraction of oxygen in the gas phaseand by referring to the 20-percent continuous volume of reaction mediumhaving the most depleted mole fraction of oxygen in the gas phase. Inthe gas phase of the distinct 20-percent continuous volume of thereaction medium containing the highest concentration of oxygen, thetime-averaged and volume-averaged oxygen concentration, on a wet basis,can be in the range of from about 3 to about 18 mole percent, in therange of from about 3.5 to about 14 mole percent, or in the range offrom 4 to 10 mole percent. In the gas phase of the distinct 20-percentcontinuous volume of the reaction medium containing the lowestconcentration of oxygen, the time-averaged and volume-averaged oxygenconcentration, on a wet basis, can be in the range of from about 0.3 toabout 5 mole percent, in the range of from about 0.6 to about 4 molepercent, or in the range of from 0.9 to 3 mole percent. Furthermore, theratio of the time-averaged and volume-averaged oxygen concentration, ona wet basis, in the most abundant 20-percent continuous volume ofreaction medium compared to the most depleted 20-percent continuousvolume of reaction medium can be in the range of from about 1.5:1 toabout 20:1, in the range of from about 2:1 to about 12:1, or in therange of from 3:1 to 9:1.

The staging of oxygen consumption rate in the reaction medium can bequantified in terms of an oxygen-STR, initially described above.Oxygen-STR was previously describe in a global sense (i.e., from theperspective of the average oxygen-STR of the entire reaction medium);however, oxygen-STR may also be considered in a local sense (i.e., aportion of the reaction medium) in order to quantify staging of theoxygen consumption rate throughout the reaction medium.

The inventors have discovered that it can be useful to cause theoxygen-STR to vary throughout the reaction medium in general harmonywith the desirable gradients disclosed herein relating to pressure inthe reaction medium and to the mole fraction of molecular oxygen in thegas phase of the reaction medium. Thus, in various embodiments, theratio of the oxygen-STR of a first distinct 20-percent continuous volumeof the reaction medium compared to the oxygen-STR of a second distinct20-percent continuous volume of the reaction medium can be in the rangeof from about 1.5:1 to about 20:1, in the range of from about 2:1 toabout 12:1, or in the range of from 3:1 to 9:1. In one embodiment, the“first distinct 20-percent continuous volume” can be located closer thanthe “second distinct 20-percent continuous volume” to the location wheremolecular oxygen is initially introduced into the reaction medium. Theselarge gradients in oxygen-STR may be desirable whether the partialoxidation reaction medium is contained in a bubble column oxidationreactor or in any other type of reaction vessel in which gradients arecreated in pressure and/or mole fraction of molecular oxygen in the gasphase of the reaction medium (e.g., in a mechanically agitated vesselhaving multiple, vertically disposed stirring zones achieved by usingmultiple impellers having strong radial flow, possibly augmented bygenerally horizontal baffle assemblies, with oxidant flow risinggenerally upwards from a feed near the lower portion of the reactionvessel, notwithstanding that considerable back-mixing of oxidant flowmay occur within each vertically disposed stiffing zone and that someback-mixing of oxidant flow may occur between adjacent verticallydisposed stirring zones). That is, when a gradient exists in thepressure and/or mole fraction of molecular oxygen in the gas phase ofthe reaction medium, the inventors have discovered that it may bedesirable to create a similar gradient in the chemical demand fordissolved oxygen.

One way of causing the local oxygen-STR to vary is by controlling thelocations of feeding the oxidizable compound and by controlling themixing of the liquid phase of the reaction medium to control gradientsin concentration of oxidizable compound according to other disclosuresherein. Other useful means of causing the local oxygen-STR to varyinclude causing variation in reaction activity by causing localtemperature variation and by changing the local mixture of catalyst andsolvent components (e.g., by introducing an additional gas to causeevaporative cooling in a particular portion of the reaction mediumand/or by adding a solvent stream containing a higher amount of water todecrease activity in a particular portion of the reaction medium).

Referring now to FIG. 6, a process is illustrated for producing purifiedterephthalic acid (“PTA”) employing an oxidation reactor system 200comprising a primary oxidation reactor 200 a and a secondary oxidationreactor 200 b. In the configuration illustrated in FIG. 6, an initialslurry can be produced from primary oxidation reactor 200 a and canthereafter be subjected to purification in a purification system 202, ofwhich secondary oxidation reactor 200 b is a part. The initial slurrywithdrawn from primary oxidation reactor 200 a can comprise solid crudeterephthalic acid (“CTA”) particles and a liquid mother liquor.Typically, the initial slurry can contain in the range of from about 10to about 50 weight percent solid CTA particles, with the balance beingliquid mother liquor. The solid CTA particles present in the initialslurry withdrawn from primary oxidation reactor 200 a can contain atleast about 400 ppmw of 4-carboxybenzaldehyde (“4-CBA”), at least about800 ppmw of 4-CBA, or in the range of from 1,000 to 15,000 ppmw of4-CBA.

Purification system 202 receives the initial slurry withdrawn fromprimary oxidation reactor 200 a and reduces the concentration of 4-CBAand other impurities present in the CTA. A purer/purified slurry can beproduced from purification system 202 and can be subjected to separationand drying in a separation system 204 to thereby produce purer solidterephthalic acid particles comprising less than about 400 ppmw of4-CBA, less than about 250 ppmw of 4-CBA, or in the range of from 10 to200 ppmw of 4-CBA.

Purification system 202 includes secondary oxidation reactor 200 b, adigester 206, and a single crystallizer 208. In secondary oxidationreactor 200 b, the initial slurry is subjected to oxidation atconditions such as described above with reference to secondary oxidationreactor 104 of FIG. 2. The slurry exiting secondary oxidation reactor200 b is introduced into digester 206. In digester 206, a furtheroxidation reaction can be performed at slightly higher temperatures thanwere used in primary oxidation reactor 200 a.

The high surface area, small particle size, and low density of the CTAparticles produced in primary oxidation reactor 200 a can cause certainimpurities trapped in the CTA particles to become available foroxidation in digester 206 without requiring complete dissolution of theCTA particles in digester 206. Thus, the temperature in digester 206 canbe lower than many similar prior art processes. The further oxidationcarried out in digester 206 can reduce the concentration of 4-CBA in theCTA by at least 200 ppmw, at least about 400 ppmw, or in the range offrom 600 to 6,000 ppmw. The digestion temperature in digester 206 can beat least about 10° C. higher than the primary oxidation temperature inreactor 200 a, about 20 to about 80° C. higher than the primaryoxidation temperature in reactor 200 a, or 30 to 50° C. higher than theprimary oxidation temperature in reactor 200 a. The digestiontemperature can be in the range of from about 160 to about 240° C., inthe range of from about 180 to about 220° C., or in the range of from190 to 210° C. In various embodiments, the purified product fromdigester 206 needs only a single crystallization step in crystallizer208 prior to separation in separation system 204. Suitable secondaryoxidation/digestion techniques are discussed in further detail in U.S.Pat. No. 7,132,566, the entire disclosure of which is expresslyincorporated herein by reference.

Terephthalic acid (e.g., PTA) produced by the system illustrated in FIG.6 can be formed of PTA particles having a mean particle size of at leastabout 40 micrometers (μm), in the range of from about 50 to about 2,000μm, or in the range of from 60 to 200 μm. The PTA particles can have anaverage BET surface area less than about 0.25 m²/g, in the range of fromabout 0.005 to about 0.2 m²/g, or in the range of from 0.01 to 0.18m²/g. PTA produced by the system illustrated in FIG. 6 is suitable foruse as a feedstock in the making of PET. Typically, PET is made viaesterification of terephthalic acid with ethylene glycol, followed bypolycondensation. In various embodiments, terephthalic acid produced byan embodiment of the present invention can be employed as a feed to thepipe reactor PET process described in U.S. Pat. No. 6,861,494, theentire disclosure of which is incorporated herein by reference.

CTA particles with the morphology disclosed herein may be particularlyuseful in the above-described oxidative digestion process for reductionof 4-CBA content. In addition, these CTA particles may provideadvantages in a wide range of other post-processes involving dissolutionand/or chemical reaction of the particles. These additionalpost-processes include, but are not limited too, reaction with at leastone hydroxyl-containing compound to form ester compounds, especially thereaction of CTA with methanol to form dimethyl terephthalate andimpurity esters; reaction with at least one diol to form ester monomerand/or polymer compounds, especially the reaction of CTA with ethyleneglycol to form polyethylene terephthalate (PET); and full or partialdissolution in solvents, including, but not limited too, water, aceticacid, and N-methyl-2-pyrrolidone, which may include further processing,including, but not limited too, reprecipitation of a more pureterephthalic acid and/or selective chemical reduction of carbonyl groupsother than carboxylic acid groups. Notably included is the substantialdissolution of CTA in a solvent comprising water coupled with partialhydrogenation that reduces the amount of aldehydes, especially 4-CBA,fluorenones, phenones, and/or anthraquinones.

Definitions

It should be understood that the following is not intended to be anexclusive list of defined terms. Other definitions may be provided inthe foregoing description, such as, for example, when accompanying theuse of a defined term in context.

As used herein, the terms “a,” “an,” and “the” mean one or more.

As used herein, the term “and/or,” when used in a list of two or moreitems, means that any one of the listed items can be employed by itselfor any combination of two or more of the listed items can be employed.For example, if a composition is described as containing components A,B, and/or C, the composition can contain A alone; B alone; C alone; Aand B in combination; A and C in combination, B and C in combination; orA, B, and C in combination.

As used herein, the terms “comprising,” “comprises,” and “comprise” areopen-ended transition terms used to transition from a subject recitedbefore the term to one or more elements recited after the term, wherethe element or elements listed after the transition term are notnecessarily the only elements that make up the subject.

As used herein, the terms “having,” “has,” and “have” have the sameopen-ended meaning as “comprising,” “comprises,” and “comprise” providedabove.

As used herein, the terms “including,” “includes,” and “include” havethe same open-ended meaning as “comprising,” “comprises,” and “comprise”provided above.

Numerical Ranges

The present description uses numerical ranges to quantify certainparameters relating to the invention. It should be understood that whennumerical ranges are provided, such ranges are to be construed asproviding literal support for claim limitations that only recite thelower value of the range as well as claim limitations that only recitethe upper value of the range. For example, a disclosed numerical rangeof 10 to 100 provides literal support for a claim reciting “greater than10” (with no upper bounds) and a claim reciting “less than 100” (with nolower bounds).

The present description uses specific numerical values to quantifycertain parameters relating to the invention, where the specificnumerical values are not expressly part of a numerical range. It shouldbe understood that each specific numerical value provided herein is tobe construed as providing literal support for a broad, intermediate, andnarrow range. The broad range associated with each specific numericalvalue is the numerical value plus and minus 60 percent of the numericalvalue, rounded to two significant digits. The intermediate rangeassociated with each specific numerical value is the numerical valueplus and minus 30 percent of the numerical value, rounded to twosignificant digits. The narrow range associated with each specificnumerical value is the numerical value plus and minus 15 percent of thenumerical value, rounded to two significant digits. For example, if thespecification describes a specific temperature of 62° F., such adescription provides literal support for a broad numerical range of 25°F. to 99° F. (62° F.+/−37° F.), an intermediate numerical range of 43°F. to 81° F. (62° F.+/−19° F.), and a narrow numerical range of 53° F.to 71° F. (62° F.+/−9° F.). These broad, intermediate, and narrownumerical ranges should be applied not only to the specific values, butshould also be applied to differences between these specific values.Thus, if the specification describes a first pressure of 110 psia and asecond pressure of 48 psia (a difference of 62 psi), the broad,intermediate, and narrow ranges for the pressure difference betweenthese two streams would be 25 to 99 psi, 43 to 81 psi, and 53 to 71 psi,respectively.

Claims not Limited to Disclosed Embodiments

The forms of the invention described above are to be used asillustration only, and should not be used in a limiting sense tointerpret the scope of the present invention. Modifications to theexemplary embodiments, set forth above, could be readily made by thoseskilled in the art without departing from the spirit of the presentinvention.

We claim:
 1. A system for producing a polycarboxylic acid by contactinga slurry with a gas-phase oxidant, said system comprising: a primaryoxidation reactor comprising a first slurry outlet; and a secondaryoxidation reactor comprising a slurry inlet, a second slurry outlet, anormally lower oxidant inlet, and a normally upper oxidant inlet,wherein said slurry inlet is in downstream fluid-flow communication withsaid first slurry outlet, wherein said secondary oxidation reactordefines therein a secondary reaction zone having a maximum length L_(s)and a maximum diameter D_(s), wherein said normally lower oxidant inletis spaced from the bottom of said secondary reaction zone by less than0.5 L_(s), wherein said normally upper oxidant inlet is spaced from thebottom of said secondary reaction zone by at least 0.5 L_(s), whereinsaid slurry inlet is spaced from the bottom of said secondary reactionzone by a distance in the range of from about 0.3 L_(s) to about 0.9L_(s).
 2. The system of claim 1, wherein said normally upper oxidantinlet and said normally lower oxidant inlet define between them a totalopen area for introducing said gas-phase oxidant into said secondaryreaction zone, wherein said normally upper oxidant inlet defines in therange of from about 5 to about 49 percent of said total open area. 3.The system of claim 1, wherein said normally upper oxidant inlet isspaced from the bottom of said secondary reaction zone by at least 0.7L_(s).
 4. The system of claim 1, wherein said normally upper oxidantinlet comprises a sparger, wherein said sparger comprises a plurality ofoxidant discharge openings.
 5. The system of claim 4, wherein a majorityof said oxidant discharge openings are oriented to discharge saidgas-phase oxidant in a normally downward direction.
 6. The system ofclaim 1, wherein said normally upper oxidant inlet is spaced less than0.4 L_(s) from said slurry inlet.
 7. The system of claim 1, wherein saidsecondary oxidation reactor comprises at least two normally upperoxidant inlets, each individually spaced from the bottom of saidsecondary reaction zone by at least 0.5 L_(s).
 8. The system of claim 1,wherein said slurry inlet is spaced from the bottom of said secondaryreaction zone by a distance in the range of from about 0.5 L_(s) toabout 0.8 L_(s), wherein said slurry inlet is spaced from the bottom ofsaid secondary reaction zone by a distance in the range of from about 9D_(s) to about 15 D_(s).
 9. The system of claim 1, wherein said reactionzone has a L_(s): D_(s) ratio in the range of from about 14:1 to about28:1.
 10. The system of claim 1, wherein said primary oxidation reactoris a bubble column reactor, wherein said secondary oxidation reactor isa bubble column reactor, wherein said primary oxidation reactor definestherein a primary reaction zone, wherein a volume ratio of said primaryreaction zone to said secondary reaction zone is in the range of fromabout 4:1 to about 50:1.
 11. A system for producing a polycarboxylicacid by contacting a slurry produced via oxidation with a gas-phaseoxidant, said system comprising: a primary oxidation reactor comprisinga first slurry outlet; and a secondary oxidation reactor comprising aslurry inlet, a second slurry outlet, a normally lower oxidant inlet anda normally upper oxidant inlet, wherein said slurry inlet is indownstream fluid-flow communication with said first slurry outlet,wherein said secondary oxidation reactor defines therein a secondaryreaction zone having a maximum length L_(s) and a maximum diameterD_(s), wherein said slurry inlet is spaced from the bottom of saidsecondary reaction zone by a distance in the range of from about 0.3L_(s) to about 0.9 L_(s), wherein said normally upper oxidant inlet isspaced above said slurry inlet by less than 0.4 L_(s).
 12. The system ofclaim 11, wherein said normally lower oxidant inlet is spaced from thebottom of said secondary reaction zone by less than 0.3 L_(s).
 13. Thesystem of claim 11, wherein said normally upper oxidant inlet is spacedfrom the bottom of said secondary reaction zone by at least 0.7 L_(s).14. The system of claim 11, wherein said normally upper oxidant inletcomprises a sparger, wherein said sparger comprises a plurality ofoxidant discharge openings.
 15. The system of claim 14, wherein amajority of said oxidant discharge openings are oriented to dischargesaid gas-phase oxidant in a normally downward direction.
 16. The systemof claim 11, wherein said slurry inlet is spaced from the bottom of saidsecondary reaction zone by a distance in the range of from about 0.5L_(s) to about 0.8 L_(s), wherein said slurry inlet is spaced from thebottom of said secondary reaction zone by a distance in the range offrom about 9 D_(s) to about 15 D_(s).
 17. The system of claim 11,wherein said reaction zone has a L_(s): D_(s) ratio in the range of fromabout 14:1 to about 28:1.
 18. The system of claim 11, wherein said firstslurry outlet and said slurry inlet have substantially the same verticalelevation, wherein said first slurry outlet and said slurry inlet arecoupled in fluid-flow communication via a substantially straight andsubstantially horizontal conduit.
 19. The system of claim 11, whereinsaid primary oxidation reactor is a bubble column reactor, wherein saidsecondary oxidation reactor is a bubble column reactor.
 20. The systemof claim 11, wherein said primary oxidation reactor defines therein aprimary reaction zone, wherein a volume ratio of said primary reactionzone to said secondary reaction zone is in the range of from about 4:1to about 50:1.
 21. A method for making a polycarboxylic acidcomposition, said method comprising: (a) subjecting a first multi-phasereaction medium comprising an oxidizable compound to oxidation in aprimary reaction zone defined in a primary oxidation reactor to therebyproduce a first slurry; and (b) contacting at least a portion of saidfirst slurry with a gas-phase oxidant in a secondary reaction zonedefined in a secondary oxidation reactor to thereby produce a secondslurry, wherein said secondary reaction zone has a maximum length L_(s)and a maximum diameter Ds, wherein a first portion of said gas-phaseoxidant is introduced into said secondary reaction zone at a firstoxidant inlet region spaced from the bottom of said secondary reactionzone by at least 0.5 L_(s), wherein said first portion of said gas-phaseoxidant constitutes in the range of from about 5 to about 49 percent ofthe total volume of said gas-phase oxidant introduced into saidsecondary reaction zone, wherein at least a portion of said first slurryis introduced into said secondary reaction zone at a slurry inlet regionspaced from the bottom of said secondary reaction zone by a distance inthe range of from about 0.3 L_(s) to about 0.9 L_(s), wherein a secondportion of said gas-phase oxidant is introduced into said secondaryreaction zone at a second oxidant inlet region spaced from the bottom ofsaid secondary reaction zone by less than 0.3 L_(s).
 22. The method ofclaim 21, wherein said first portion of said gas-phase oxidantconstitutes in the range of from about 5 to about 35 percent of thetotal volume of said gas-phase oxidant introduced into said secondaryreaction zone, wherein said first oxidant inlet region is spaced fromthe bottom of said secondary reaction zone by at least 0.7 L_(s). 23.The method of claim 21, wherein said first oxidant inlet region iswithin 0.4 L_(s) of said slurry inlet region.
 24. The method of claim21, wherein at least a portion of said gas-phase oxidant and at least aportion of said first slurry combine in said secondary reaction zone toform a second multi-phase reaction medium, wherein when the entirevolume of said second multi-phase reaction medium is theoreticallypartitioned into 20 discrete horizontal slices of equal volume, no twoadjacent horizontal slices have a combined time-averaged andvolume-averaged oxygen content of less than 7 parts per million byweight (“ppmw”).
 25. The method of claim 24, wherein none of saidhorizontal slices has a time-averaged and volume-averaged oxygen contentof less than 7 ppmw.
 26. The method of claim 21, wherein at least 75weight percent of said first slurry is introduced into said secondaryreaction zone at said slurry inlet region, wherein said slurry inletregion is spaced from the bottom of said secondary reaction zone by adistance in the range of from about 0.5 L_(s) to about 0.8 L_(s). 27.The method of claim 21, wherein said first and second slurries eachcomprise para-toluic acid in the liquid phase, wherein said secondslurry has a time-averaged and volume-averaged concentration ofliquid-phase para-toluic acid that is less than 50 percent of thetime-averaged and volume-averaged concentration of liquid-phasepara-toluic acid in said first slurry.
 28. The method of claim 21,wherein said oxidizable compound is para-xylene, wherein saidpolycarboxylic acid is terephthalic acid, wherein said gas-phase oxidantis air.
 29. The method of claim 21, wherein said primary oxidationreactor is a bubble column reactor, wherein said secondary oxidationreactor is a bubble column reactor.
 30. The method of claim 21, whereinsaid secondary reaction zone has an L_(s): D_(s) ratio in the range offrom about 14:1 to about 28:1.